CN110785397B

CN110785397B - Process for producing acetic acid

Info

Publication number: CN110785397B
Application number: CN201880006880.XA
Authority: CN
Inventors: 清水雅彦
Original assignee: Daicel Corp
Current assignee: Daicel Corp
Priority date: 2018-05-29
Filing date: 2018-05-29
Publication date: 2022-12-02
Anticipated expiration: 2038-05-29
Also published as: JPWO2019229858A1; EP3792244B1; WO2019229858A1; CN110785397A; JP6626987B1; US10737996B2; TWI789521B; AR114478A1; US20200123090A1; EP3792244A4; TW202003438A; EP3792244A1

Abstract

The present invention provides a method for industrially efficiently producing acetic acid having a good potassium permanganate test value without incurring a high cost. In the acetic acid production method, the acetaldehyde separation and removal step includes a step of separating and removing acetaldehyde as a by-product in an industrially advantageous manner and distilling an aqueous phase and/or an organic phase of a condensate at the top of the low-boiling-point removal column in a crotonaldehyde removal column, and the crotonaldehyde removal column is operated so as to satisfy at least one of the following conditions (a-i) to (a-iii) while setting a reflux ratio of the low-boiling-point removal column to 2 or more (in the case of refluxing the aqueous phase): (a-i) adjusting the reflux ratio of the distillation column to 0.01 or more; (a-ii) making the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate to the crotonaldehyde concentration (mass ppm) in the feed liquid of the distillation column lower than 1; (a-iii) making the ratio of the crotonaldehyde concentration (mass ppm) in the bottoms of the distillation column to the crotonaldehyde concentration (mass ppm) in the feed liquid (the former/the latter) greater than 1.

Description

Process for producing acetic acid

Technical Field

The present invention relates to a process for the preparation of acetic acid.

Background

As an industrial production method of acetic acid, a methanol carbonylation process is known. In this process, for example, acetic acid is produced by reacting methanol and carbon monoxide in the presence of a catalyst in a reaction tank, the resulting reaction mixture is separated in an evaporation tank into a vapor phase containing acetic acid and low-boiling components and a residual liquid phase containing acetic acid and catalyst, and the vapor phase is distilled by a distillation column (low-boiling column) to be separated into an overhead stream containing low-boiling components and an acetic acid stream, which is further purified to obtain product acetic acid. In this process, acetaldehyde is by-produced during the reaction, and this acetaldehyde causes a reduction in the quality of acetic acid as a product.

In addition, acetaldehyde formed by the reduction of methyl iodide is converted to crotonaldehyde by aldol condensation, which results in deterioration of the potassium permanganate test value (permanganate time) of the acetic acid product. Further, crotonaldehyde is converted to 2-ethyl crotonaldehyde by aldol condensation with acetaldehyde, and 2-ethyl crotonaldehyde also causes deterioration in the potassium permanganate test value of the acetic acid product. However, crotonaldehyde is more deteriorated in potassium permanganate test value per unit mass than 2-ethyl crotonaldehyde, and when crotonaldehyde is contained in product acetic acid, deterioration in quality becomes more remarkable.

Conventionally, in order to reduce acetaldehyde, crotonaldehyde, and 2-ethylcrotonaldehyde, the following two methods have been generally industrially adopted: (i) A method for suppressing the formation of crotonaldehyde in a reaction system by separating a condensate of an overhead stream of the above-mentioned low-boiling column into an aqueous phase and an organic phase by a decanter, distilling the aqueous phase in the above-mentioned column by an acetaldehyde removing column, extracting the condensate of the overhead stream (containing acetaldehyde and methyl iodide) with water to separate and remove acetaldehyde, and reducing acetaldehyde in methyl iodide which is recycled to the reaction system; and (ii) a method of oxidatively decomposing crotonaldehyde contained in crude acetic acid obtained in the course of the purification step directly using ozone (patent documents 1 and 2). However, both the acetaldehyde separation and removal equipment and the ozone treatment equipment are expensive. Conventionally, these methods have been relied upon entirely to improve the potassium permanganate test value of the product acetic acid, which leads to an increase in equipment costs.

In addition, it is known that in the methanol acetic acid process, alkanes are produced as impurities. The alkane is an impurity having a carbon number of 3 or more and a boiling point higher than methyl iodide and methyl acetate. Mainly saturated or unsaturated hydrocarbons, but may contain oxygen atoms or iodine atoms in the molecule. The following technique is disclosed in japanese patent laid-open No. 4-295445: in order to remove the paraffins, the organic phase in the overhead condensate of the low-boiling column is distilled by a distillation column (dealkanizing column), the overhead distillate containing methyl iodide, methyl acetate and carbonyl impurities is recycled to the reactor or supplied to an aldehyde removal column, and water is added to the bottom column bottom residue containing paraffins, water and acetic acid to extract them, the aqueous phase containing acetic acid is recycled to the reactor, and the organic phase containing paraffins is discarded. However, there is no disclosure or suggestion in this document of how to improve the potassium permanganate test value of the product acetic acid.

Documents of the prior art

Patent document

Patent document 1: japanese laid-open patent publication No. H07-25813

Patent document 2: japanese Kokai publication Hei-2001-508405

Patent document 3: japanese laid-open patent publication No. 4-295445

Disclosure of Invention

Problems to be solved by the invention

The inventors of the present invention have studied and found that in the method of distilling the aqueous phase of the two phases separated by the decanter of the low boiling point removal column by the acetaldehyde removal column as described in patent document 2, the separation efficiency tends to be generally higher because the concentration of acetaldehyde is higher in the aqueous phase than in the organic phase. However, in this method, since a liquid containing a large amount of water having a large latent heat of vaporization is distilled, not only the processing energy is increased, but also a distillation apparatus is easily corroded by hydrogen iodide generated by hydrolysis of methyl iodide contained in a small amount in an aqueous phase, and thus a distillation apparatus of an expensive material having extremely high corrosion resistance is required. Since a considerable amount of methyl iodide is present in the aqueous phase, the number of distillation trays is required to be about the same as that in the case of distilling the organic phase (methyl iodide phase) in the acetaldehyde removal column described later in order to separate methyl iodide and acetaldehyde having a small difference in boiling points in addition to water.

On the other hand, a method is also known in which the organic phase of the two phases separated by the decanter of the low-boiling-point removal column is distilled by an acetaldehyde removal column, and the condensate of the overhead stream is extracted with water to separate and remove acetaldehyde. In this method, there are the following advantages: the main component of the organic phase is methyl iodide having a small latent heat of vaporization and therefore the processing energy is small, and the problem of corrosion is less likely to occur as compared with the case of processing the aqueous phase. However, as described above, the inventors of the present invention have studied to find the following disadvantages: generally, the separation efficiency is low because of the lower acetaldehyde concentration in the organic phase compared to the aqueous phase. Thus, both the method of subjecting the aqueous phase to the acetaldehyde removal treatment and the method of subjecting the organic phase to the acetaldehyde removal treatment have advantages and disadvantages.

Accordingly, an object of the present invention is to provide a method for industrially efficiently producing acetic acid having a good potassium permanganate test value without incurring high costs.

Further, an object of the present invention is to provide a method capable of separating and removing acetaldehyde as a by-product in an industrially advantageous manner by examining advantages and disadvantages of a method for subjecting an aqueous phase to acetaldehyde-removing treatment and a method for subjecting an organic phase to acetaldehyde-removing treatment in a liquid separation step of separating a process stream in a methanol carbonylation process into an aqueous phase and an organic phase.

Means for solving the problems

The present inventors have found that the crotonaldehyde concentration in the bottom residue of a dehydration column can be reduced and the potassium permanganate test value of the bottom residue can be greatly improved by setting the reflux ratio of a low-boiling column to a specific value or more and providing a distillation column (crotonaldehyde removal column) for treating the organic phase of the top condensate of the low-boiling column and specifying the operating conditions of the crotonaldehyde removal column. More specifically, crotonaldehyde has a boiling point (104 ℃) lower than that of acetic acid (117 ℃), and thus crotonaldehyde can be concentrated to the top of the distillation column by increasing the reflux ratio of the low-boiling-point removal column. If the concentrated crotonaldehyde is recycled to the reaction tank, it reacts with acetaldehyde in the reaction tank to form 2-ethyl crotonaldehyde. Further, crotonaldehyde reacts with hydrogen in the reaction tank to produce butanol, and this butanol reacts with acetic acid to produce butyl acetate. Compared with crotonaldehyde, 2-ethyl crotonaldehyde has small influence on the potassium permanganate test value, and butanol and butyl acetate do not influence the potassium permanganate test value at all and are harmless. Therefore, the quality of acetic acid tends to be further improved. Since 2-ethylcrotonaldehyde and butyl acetate have boiling points of 137 ℃ and 126 ℃ and higher than the boiling point (117 ℃) of acetic acid, respectively, if the reflux ratio of the low-boiling-point removal column is increased, the overhead concentration of these components is further reduced, and the bottom liquid from the low-boiling-point removal column is recycled to the reaction system and concentrated, or a part of the side stream higher than the feed position is sent to a subsequent step and contained in the acetic acid as a product. On the other hand, by subjecting the condensate on the top of the low-boiling-point-removal column, in which crotonaldehyde is concentrated, to a distillation treatment different from the acetaldehyde removal treatment, it is possible to efficiently separate useful methyl iodide from unnecessary crotonaldehyde. That is, if the organic phase in the condensate on the top of the low-boiling column is subjected to distillation, methyl iodide can be obtained as a distillate on the top of the column together with methyl acetate. It can be recycled to a decanter or a reaction tank for storing condensate at the top of the low-boiling-point removal tower. Further, crotonaldehyde can be obtained as a bottom residue together with other high boiling impurities (e.g., 2-ethylcrotonaldehyde, butyl acetate, and alkanes) and acetic acid. The residue liquid in the tower bottom is removed out of the system and is discarded. As for water, it may be concentrated to the top of the column or may be extracted from the bottom of the column. It should be noted that a conventionally known dealkanization column may be used as the crotonaldehyde removing column. The dealkanizing column may be operated constantly, but may be operated intermittently when the generation rate of the paraffin is low. In the case of the batch operation, the potassium permanganate test value of the product acetic acid is lowered, and therefore, it is necessary to maintain the product quality by changing the acetaldehyde removal treatment, the ozone treatment, the operation conditions, and the like. The test value of potassium permanganate can be improved by discharging 2-ethylcrotonaldehyde, which is present in the very small amount at the top of the low-boiling column, to the outside of the system by the above operation, similarly to crotonaldehyde, but the effect is limited because 2-ethylcrotonaldehyde having a high boiling point is hardly concentrated at the top of the low-boiling column. The organic phase in the condensate on the top of the low-boiling column is mainly supplied to the crotonaldehyde removing column, but the aqueous phase in the condensate on the top of the low-boiling column may be supplied to the crotonaldehyde removing column in addition to or instead of the organic phase. Thus, the potassium permanganate test value of the product acetic acid can be easily improved, and hence the acetaldehyde removal facility and the ozone treatment facility can be eliminated or the scale can be reduced, and the steam and electricity charges can be reduced.

In addition, as a result of intensive studies to achieve the above object, the present inventors have conducted a liquid-liquid equilibrium experiment for designing a condensate composition of a top stream of a low-boiling column, focusing on distribution ratios of acetaldehyde (acetaldehyde concentration in an aqueous phase/acetaldehyde concentration in an organic phase) under various conditions in a liquid separation step. As a result, it was found that there was a certain correlation between the acetaldehyde concentration in the aqueous phase and the acetaldehyde concentration in the organic phase obtained by separation with a decanter of a low-boiling point removal column, the liquid temperature at the time of separation, and the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase, and the distribution ratio of acetaldehyde. As described above, since the method of subjecting the aqueous phase to the acetaldehyde-removing treatment and the method of subjecting the organic phase to the acetaldehyde-removing treatment have advantages and disadvantages, respectively, it is industrially advantageous to adopt the former method when the acetaldehyde distribution ratio is high to a certain degree (when acetaldehyde is distributed in a large amount in the aqueous phase), the latter method when the acetaldehyde distribution ratio is low to a certain degree (when acetaldehyde is distributed in a large amount in the organic phase), and the both methods in the intermediate region. In particular, when a method of removing acetaldehyde from an aqueous phase is employed, since expensive equipment and facilities having high corrosion resistance are required, the acetaldehyde distribution rate must be strictly recognized.

The present invention has been completed based on these findings and through further studies.

Namely, the present invention provides a method for producing acetic acid, which comprises:

a carbonylation reaction step in which methanol and carbon monoxide are reacted in a reaction vessel in the presence of a catalyst system comprising a metal catalyst and methyl iodide, and acetic acid, methyl acetate, and water to produce acetic acid;

a separation step of separating and obtaining a stream containing a metal catalyst, an acetic acid stream rich in acetic acid, and a stream rich in low-boiling components compared with the acetic acid stream from the reaction mixture obtained in the carbonylation reaction step, using 1 or more evaporation tanks and/or distillation columns;

a liquid separation step of separating a process stream containing at least water, acetic acid, methyl iodide, and acetaldehyde into an aqueous phase and an organic phase;

an acetaldehyde separation/removal step of separating and removing acetaldehyde from the process stream, wherein a treatment target is selected from at least a part of the aqueous phase and/or the organic phase and treated based on a liquid temperature at the time of liquid separation in the liquid separation step, an acetaldehyde concentration, an acetaldehyde distribution ratio, a methyl acetate concentration, or a methyl acetate distribution ratio of the aqueous phase and/or the organic phase;

A recycling process of recycling at least a portion of the aqueous phase and/or the organic phase to the reaction tank; and

a crotonaldehyde removal step of treating at least another part of the aqueous phase and/or the organic phase with a distillation column to separate and remove crotonaldehyde,

in the manufacturing method of the present invention, the substrate is,

the separation step comprises a step of separating a 1 st acetic acid stream enriched in acetic acid and a top stream enriched in low-boiling components compared with the 1 st acetic acid stream in a 1 st distillation column, wherein the reflux ratio of the 1 st distillation column is such that the reflux ratio of the aqueous phase is 2 or more when only the aqueous phase is refluxed, the reflux ratio of the organic phase is 1 or more when only the organic phase is refluxed, and the total reflux ratio of the aqueous phase and the organic phase is 1.5 or more when both the aqueous phase and the organic phase are refluxed, and wherein the distillation column is operated so as to satisfy at least 1 of the following conditions (a-i) to (a-iii) in the crotonaldehyde removing step:

(a-i) adjusting the reflux ratio of the distillation column to 0.01 or more;

(a-ii) making the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate to the crotonaldehyde concentration (mass ppm) in the feed liquid of the distillation column lower than 1;

(a-iii) making the ratio of the crotonaldehyde concentration (mass ppm) in the bottoms of the distillation column to the crotonaldehyde concentration (mass ppm) in the feed liquid (the former/the latter) greater than 1.

In the method for producing acetic acid according to the present invention, it is preferable that the liquid separation step satisfies at least one of the following conditions (b-i) to (b-v), and at least a part of the aqueous phase is treated in the acetaldehyde separation and removal step,

(b-i) the acetaldehyde concentration in the aqueous phase is 28.1 mass% or less, and/or the acetaldehyde concentration in the organic phase is 24.8 mass% or less;

(b-ii) the temperature at the time of liquid separation is 70 ℃ or lower;

(b-iii) the methyl acetate concentration in the aqueous phase is 12.0 mass% or less, and/or the methyl acetate concentration in the organic phase is 47.6 mass% or less, and/or the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase is 59.6 mass% or less;

(b-iv) an acetaldehyde partition rate [ { acetaldehyde concentration (mass%) of the aqueous phase }/{ acetaldehyde concentration (mass%) of the organic phase } is 1.1 or more; and

(b-v) the partition rate of methyl acetate [ { methyl acetate concentration (mass%) of the aqueous phase)/(methyl acetate concentration (mass%) of the organic phase } is 0.25 or more.

In the method for producing acetic acid according to the present invention, it is preferable that the liquid separation step satisfies all of the conditions (b-i), (b-ii), (b-iii), (b-iv), and (b-v).

In the method for producing acetic acid according to the present invention, it is preferable that the liquid separation step satisfies at least one of the following conditions (b '-i) to (b' -v), and at least a part of the organic phase is treated in the acetaldehyde separation and removal step,

(b' -i) the acetaldehyde concentration in the aqueous phase is 0.045% by mass or more, and/or the acetaldehyde concentration in the organic phase is 0.013% by mass or more;

(b' -ii) the temperature for liquid separation is-5 ℃ or higher;

(b' -iii) the methyl acetate concentration in the aqueous phase is 1.2% by mass or more, and/or the methyl acetate concentration in the organic phase is 2.2% by mass or more, and/or the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase is 3.4% by mass or more;

(b' -iv) an acetaldehyde partition rate [ { acetaldehyde concentration (mass%) of the aqueous phase }/{ acetaldehyde concentration (mass%) of the organic phase } is 4.1 or less; and

(b' -v) the partition rate of methyl acetate [ { methyl acetate concentration (mass%) of the aqueous phase)/(methyl acetate concentration (mass%) of the organic phase } is 0.8 or less.

In the method for producing acetic acid of the present invention, the liquid separation step preferably satisfies all of the conditions (b ' -i), (b ' -ii), (b ' -iii), (b ' -iv), and (b ' -v).

In the method for producing acetic acid according to the present invention, the separation step preferably includes:

an evaporation step of separating the reaction mixture obtained in the carbonylation reaction step into a vapor stream and a residual liquid stream in an evaporation tank; and

and a low-boiling point removal step wherein the vapor stream is separated in a 1 st distillation column into a 1 st overhead stream enriched in at least one low-boiling component selected from methyl iodide and acetaldehyde and a 1 st acetic acid stream enriched in acetic acid, and the 1 st overhead stream is condensed and separated to obtain an aqueous phase and an organic phase.

In the above acetic acid production method of the present invention, the separation step preferably further comprises: a dehydration step of separating the 1 st acetic acid stream into a 2 nd overhead stream enriched in water and a 2 nd acetic acid stream enriched in acetic acid compared to the 1 st acetic acid stream using a 2 nd distillation column.

The reflux ratio of the 2 nd distillation column is, for example, 0.1 or more.

Can be as follows: the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.98 mass ppm or less, and/or the 2-ethyl crotonaldehyde concentration is 1.0 mass ppm or less, and/or the butyl acetate concentration is 15 mass ppm or less.

The potassium permanganate test value of the 2 nd acetic acid stream described above is preferably over 50 minutes.

In the method for producing acetic acid of the present invention, there may be: the crotonaldehyde concentration in the vapor stream supplied to the 1 st distillation column is 0 to 5.0 mass ppm, and/or the 2-ethylcrotonaldehyde concentration is 0 to 3.0 mass ppm, and/or the butyl acetate concentration is 0.1 to 13.0 mass ppm.

Can be as follows: the crotonaldehyde concentration in the 1 st acetic acid stream is 1.3 mass ppm or less, and/or the 2-ethyl crotonaldehyde concentration is 1.0 mass ppm or less, and/or the butyl acetate concentration is 15 mass ppm or less.

In the acetic acid production process of the present invention, the catalytic system may further comprise an ionic iodide.

In the production process of acetic acid according to the present invention, at least a part of the raffinate after separating acetaldehyde from at least a part of the aqueous phase and/or the organic phase may be recycled to the reaction tank.

The crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is, for example, 0.01 to 50 mass ppm.

In the crotonaldehyde removal step, the distillation column is preferably operated so as to satisfy all of the conditions (a-i) to (a-iii) described above.

In the above-mentioned crotonaldehyde removing step, the distillation is preferably carried out by a batch process.

The amount of the distillation column in the crotonaldehyde removing step is, for example, 0.0001 to 50 parts by mass per 100 parts by mass of the amount of the vapor stream supplied to the 1 st distillation column

In the acetic acid production process of the present invention, it may further comprise: a washing step of separating an off-gas from the process into a stream enriched in carbon monoxide and a stream enriched in acetic acid by subjecting the off-gas to an absorption treatment with an absorption solvent containing at least acetic acid.

ADVANTAGEOUS EFFECTS OF INVENTION

According to the present invention, since crotonaldehyde can be efficiently removed by the crotonaldehyde removing step while the reflux ratio of the low-boiling point removing column is set to a specific value or more, high-quality acetic acid having a good potassium permanganate test value (also referred to as "permanganate time" or "discoloration time") can be industrially efficiently produced without providing a large-scale acetaldehyde removing facility or ozone treatment facility. In addition, according to the present invention, in the acetaldehyde separation and removal step, the treatment target is selected from at least a part of the aqueous phase and/or the organic phase and treated based on the liquid temperature at the time of liquid separation in the liquid separation step, the acetaldehyde concentration, the acetaldehyde partition rate, the methyl acetate concentration in the aqueous phase and/or the organic phase, or the methyl acetate partition rate, and therefore acetaldehyde by-produced in the methanol carbonylation process can be efficiently removed in an industrially advantageous manner.

Drawings

FIG. 1 shows a flow diagram for the production of acetic acid according to one embodiment of the present invention.

FIG. 2 is a schematic flow chart showing an example of the system for separating and removing acetaldehyde.

FIG. 3 is a schematic flow diagram showing another example of the system for separating and removing acetaldehyde.

FIG. 4 is a schematic flow diagram showing another example of the acetaldehyde separation and removal system.

FIG. 5 is a schematic flow chart showing another example of the system for separating and removing acetaldehyde

FIG. 6 is a schematic view of a liquid-liquid equilibrium measuring apparatus used in an experimental example of the present invention.

Detailed Description

In the production method of acetic acid of the present invention, there is provided a production method of acetic acid, which comprises:

a separation step of separating and obtaining a stream containing a metal catalyst, an acetic acid stream enriched in acetic acid, and a stream enriched in low-boiling components compared with the acetic acid stream from the reaction mixture obtained in the carbonylation reaction step, by using 1 or more evaporation tanks and/or distillation columns;

An acetaldehyde separation and removal step of separating and removing acetaldehyde from the process stream, wherein a treatment target is selected from at least a part of the aqueous phase and/or the organic phase and treated based on a liquid temperature at the time of liquid separation in the liquid separation step, an acetaldehyde concentration, an acetaldehyde distribution ratio, a methyl acetate concentration, or a methyl acetate distribution ratio of the aqueous phase and/or the organic phase;

a recycling process for recycling at least a portion of the aqueous phase and/or organic phase to the reaction tank; and

a crotonaldehyde removing step of treating at least another part of the aqueous phase and/or the organic phase with a distillation column to separate and remove crotonaldehyde,

in the manufacturing method of the present invention, the substrate is,

(a-i) adjusting the reflux ratio of the distillation column to 0.01 or more;

In the above-described process for producing acetic acid of the present invention, the separation step preferably includes, for example: an evaporation step of separating the reaction mixture obtained in the carbonylation reaction step into a vapor stream and a residual liquid stream in an evaporation tank; a low-boiling removal step of distilling the above vapor stream to separate a stream (for example, an overhead stream, specifically, the 1 st overhead stream) having a higher concentration of low-boiling components than the 1 st acetic acid stream and the 1 st acetic acid stream having a higher concentration of acetic acid; and a dehydration step of distilling the 1 st acetic acid stream to separate it into a water-enriched overhead stream (2 nd overhead stream) and a 2 nd acetic acid stream enriched in acetic acid as compared with the 1 st acetic acid stream. The distillation column used in the low-boiling point removal step may be referred to as a 1 st distillation column (low-boiling point removal column) and the distillation column used in the dehydration step may be referred to as a 2 nd distillation column (dehydration column).

In the separation step, instead of the evaporation step and the low-boiling point removal step, a step of separating the reaction mixture obtained in the carbonylation reaction step into a stream containing the metal catalyst, a stream rich in the low-boiling point component (for example, an overhead stream), and a 1 st acetic acid stream rich in acetic acid (evaporation low-boiling point removal step) may be provided. In addition, the separation step may be a low-boiling removal step (so-called low-boiling removal dehydration step) having a function of the dehydration step in place of the low-boiling removal step and the dehydration step, in which the vapor stream is distilled and separated into a stream rich in low-boiling components (for example, an overhead stream) and an acetic acid stream dehydrated to a water concentration equivalent to that of the 2 nd acetic acid stream. Thus, the evaporation and low-boiling-point removal step may be a step (evaporation and low-boiling-point removal and dehydration step) further including the function of the dehydration step. The acetic acid stream enriched with acetic acid obtained in the low-boiling dehydration step and the evaporation low-boiling dehydration step corresponds to the 2 nd acetic acid stream. The distillation column used in the evaporation and low-boiling-point removal step, the low-boiling-point removal and dehydration step, and the evaporation and low-boiling-point removal and dehydration step corresponds to the 1 st distillation column.

The acetaldehyde separation and removal step is a step of separating and removing acetaldehyde from the process stream, and is a step of selecting and treating an object to be treated from at least a part of the aqueous phase and/or the organic phase based on a liquid temperature at the time of liquid separation in the liquid separation step, an acetaldehyde concentration, an acetaldehyde partition rate, a methyl acetate concentration, or a methyl acetate partition rate of the aqueous phase and/or the organic phase.

In the present specification, the term "process stream" refers to a step in an acetic acid production apparatus in which a process unit operation such as reaction, evaporation, distillation, cooling, condensation, liquid separation, storage, or absorption is performed, or a liquid phase or a gas phase in an apparatus or equipment for performing a process unit operation. Examples thereof include a liquid phase or a gas phase in a pipe, a reaction tank, an evaporation tank, and a distillation column.

The acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase are determined by the composition of the liquid to be supplied for liquid separation (hereinafter, may be referred to as "liquid supplied in the liquid separation step"), and the temperature at the time of liquid separation. The higher the acetaldehyde concentration in the feed liquid in the liquid separation step, the higher the acetaldehyde concentration in the aqueous phase and the organic phase; the higher the methyl acetate concentration in the feed liquid in the liquid separation step, the higher the methyl acetate concentration in the aqueous phase and the organic phase. As shown in the examples, the higher the temperature at the time of liquid separation, the higher the proportion of acetaldehyde distributed in the organic phase. Further, the acetaldehyde concentration and the methyl acetate concentration in the feed liquid in the liquid separation step can be controlled by, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporation tank, and the distillation conditions in the distillation column. In general, the higher the acetaldehyde concentration and the methyl acetate concentration in the reaction mixture, the higher the acetaldehyde concentration and the methyl acetate concentration, respectively, in the feed liquid to the liquid separation step. The following tendency was exhibited with respect to the acetaldehyde concentration in the reaction mixture: the reaction temperature, hydrogen partial pressure, methyl iodide concentration, water concentration, catalyst concentration and lithium iodide concentration in the reaction system are increased as they are higher, and the CO partial pressure and methyl acetate concentration are decreased as they are higher, respectively (see jp 2006-182691 a). Further, since methyl acetate is produced by the esterification reaction of acetic acid and methanol, the higher the concentration of acetic acid and the higher the concentration of methanol in the reaction system, respectively, the higher the concentration of methyl acetate in the reaction mixture, the higher the concentration of water, and the lower the concentration of water. In this manner, by adjusting the reaction conditions in the reaction tank and the operating conditions of the evaporation step and the distillation step performed before the liquid separation step, the composition of the liquid supplied in the liquid separation step, and further the acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase can be adjusted. The present inventors have found that there is a certain correlation between the acetaldehyde concentration in the aqueous phase and the acetaldehyde concentration in the organic phase obtained by separation using a decanter of a low-boiling point removal column, the liquid temperature at the time of separation, and the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase, and the distribution ratio of acetaldehyde. As described above, since the method of subjecting the aqueous phase to the acetaldehyde-removing treatment and the method of subjecting the organic phase to the acetaldehyde-removing treatment have advantages and disadvantages, respectively, it is industrially advantageous to adopt the former method when the acetaldehyde distribution ratio is high to a certain degree (when acetaldehyde is distributed in a large amount in the aqueous phase), the latter method when the acetaldehyde distribution ratio is low to a certain degree (when acetaldehyde is distributed in a large amount in the organic phase), and the both methods in the intermediate region.

Further, the methyl acetate concentration of the feed to the low-boiling point removal column can be controlled by controlling the methyl acetate concentration in the reaction tank. The methyl acetate concentration in the reaction tank can be determined by, for example, floating the temperature, rhodium complex (Rh), methyl iodide (MeI), and H in the reaction tank up and down under the condition that the feed amounts of methanol and CO are constant ₂ O, hydrogen partial pressure, carbon monoxide partial pressure, lithium iodide (LiI), and other factors that affect the reaction rate. I.e. if increasing the temperature, rh, meI, H ₂ O, hydrogen partial pressure, carbon monoxide partial pressure, liI, etc., the methyl acetate concentration in the reaction tank decreases, and if it decreases, the methyl acetate concentration increases. In addition, if these conditions are made constant, that is, the temperature, rh, meI, H are made constant ₂ On the basis of certain conditions such as O, hydrogen partial pressure, carbon monoxide partial pressure, liI and the like, the methyl acetate concentration is increased when the feeding amount of methanol and CO is increased, and the methyl acetate concentration is decreased when the feeding amount of methanol and CO is decreased. Since the operation is performed under a certain condition with respect to the methyl acetate concentration in the vapor (feed to the low-boiling-point removal column) obtained by evaporating the reaction liquid in the evaporation tank, the evaporation rate changes in proportion to the methyl acetate concentration in the reaction liquid. When the evaporation tank is heated and cooled, the methyl acetate concentration also changes due to a change in the evaporation rate. For example, in the case of heating, the acetic acid concentration in the vapor of the evaporation tank increases and the methyl acetate concentration decreases. On the other hand, in the case of cooling, the opposite phenomenon occurs, that is, the acetic acid concentration in the vapor in the evaporation tank decreases and the methyl acetate concentration increases.

The recycling step may include: at least a portion of the aqueous phase and/or organic phase resulting from condensing the 1 st overhead stream, and/or a portion of the 2 nd overhead stream is recycled to the overhead stream recycling process of the reaction tank.

In the present invention, crotonaldehyde is concentrated to the top of the column by increasing the reflux ratio of the 1 st distillation column, and at least a part of the aqueous phase and/or organic phase of the condensate on the top of the lower boiling column in which crotonaldehyde is concentrated is recycled to the reaction tank. Since crotonaldehyde is concentrated to the top of the column, the concentration of crotonaldehyde in the 1 st acetic acid stream is reduced, and as a result, acetic acid which is a product having a good potassium permanganate test value can be obtained. In addition, crotonaldehyde recycled to the reaction tank is converted into 2-ethyl crotonaldehyde having little influence on the potassium permanganate test value or butyl acetate having no influence at all by the reaction of crotonaldehyde + acetaldehyde → 2-ethyl crotonaldehyde, crotonaldehyde + hydrogen → butanol, butanol + acetic acid → butyl acetate, and thus the quality of acetic acid as a product can be improved. In the present invention, at least another part of the aqueous phase and/or organic phase of the condensate on the top of the low-boiling column from which crotonaldehyde has been concentrated is treated by the crotonaldehyde removing column, and as a result, crotonaldehyde has a high boiling point of 104 ℃ and is extracted as a high boiling compound together with acetic acid and alkanes to the bottom residue side of the column and discharged to the outside of the system, whereby the potassium permanganate test value of the product acetic acid can be further improved. The overhead condensate of the crotonaldehyde removing column contains useful components (for example, methyl iodide, methyl acetate, etc.), and therefore, the overhead condensate can be recycled to a decanter or a reaction tank for storing the overhead condensate of the low boiling removing column.

Regarding the reflux ratio of the 1 st distillation column, when only the aqueous phase is refluxed in the 1 st distillation column, the reflux ratio of the aqueous phase (reflux amount of the aqueous phase/distilled amount of the aqueous phase) is 2 or more, preferably 3 or more, more preferably 5 or more, further preferably 8 or more, and particularly preferably 12 or more. In the case where only the organic phase is refluxed in the 1 st distillation column, the reflux ratio of the organic phase (reflux amount of the organic phase/distillation amount of the organic phase) is 1 or more, preferably 1.5 or more, more preferably 2 or more, further preferably 4 or more, and particularly preferably 5 or more. In addition, when both the aqueous phase and the organic phase are refluxed in the 1 st distillation column, the total reflux ratio of the aqueous phase and the organic phase (the total reflux amount of the aqueous phase and the organic phase/the total distillation amount of the aqueous phase and the organic phase) is 1.5 or more, preferably 2.3 or more, more preferably 3.5 or more, further preferably 6 or more, and particularly preferably 8.5 or more. In addition, when the aqueous phase is refluxed in the 1 st distillation column, the reflux ratio of the aqueous phase (amount of reflux of the aqueous phase/amount of distillate of the aqueous phase) is preferably 2 or more, more preferably 3 or more, further preferably 5 or more, particularly preferably 8 or more, and particularly preferably 12 or more. The upper limit of the reflux ratio of the 1 st distillation column may be, for example, 3000 (particularly 1000) or may be 100 (particularly 30) in any case.

In the above (a-i), the reflux ratio of the crotonaldehyde removing column is preferably 0.05 or more, more preferably 0.5 or more, further preferably 5 or more, and particularly preferably 20 or more (for example, 30 or more). The upper limit of the reflux ratio of the crotonaldehyde removing column is, for example, 1000. In the above (a-ii), the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate of the crotonaldehyde removing column to the crotonaldehyde concentration (mass ppm) in the feed liquid is preferably 0.95 or less, more preferably 0.80 or less, still more preferably 0.70 or less, and particularly preferably 0.60 or less (for example, 0.50 or less, particularly preferably 0.30 or less, and particularly preferably 0.20 or less). In the above (a-iii), the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the bottom residue of the crotonaldehyde rejection column to the crotonaldehyde concentration (mass ppm) in the feed liquid is preferably 1.2 or more, more preferably 1.5 or more, further preferably 2.0 or more, particularly preferably 3.0 or more (for example, 4.0 or more, particularly preferably 5.0 or more), and particularly preferably 10 or more (for example, 20 or more). When the crotonaldehyde removing column is operated so as to satisfy at least one of the conditions (a-i) to (a-iii), crotonaldehyde is condensed at the bottom of the column and discharged as a column bottom residue together with other high boiling impurities such as alkanes and acetic acid.

In the present specification, the "reflux ratio" in the distillation column means "reflux amount/distillate amount". Here, the "reflux amount" refers to the amount of a liquid returned to the distillation column in the column top liquid of the distillation column, and the "distillate amount" refers to the amount of a liquid discharged to the outside of the distillation column without being returned to the distillation column in the column top liquid of the distillation column.

In the present invention, the phase to be supplied to the acetaldehyde separation and removal step can be determined based on the acetaldehyde concentration in the aqueous phase, the acetaldehyde concentration in the organic phase, the methyl acetate concentration in the aqueous phase, the methyl acetate concentration in the organic phase, the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase, the temperature (liquid temperature) at the time of liquid separation, the acetaldehyde partition rate (AD partition rate) obtained by the following formula, or the methyl acetate partition rate (MA partition rate) obtained by the following formula in the liquid separation step (for example, liquid separation by the decanter 4) of separating the process stream containing at least water, acetic Acid (AC), methyl iodide (MeI), and Acetaldehyde (AD). From the other side, suitable conditions for separation when the organic phase is supplied to the acetaldehyde separation and removal step and suitable conditions for separation when the aqueous phase is supplied to the acetaldehyde separation and removal step are shown.

AD partition rate = { AD concentration (mass%) of aqueous phase)/(AD concentration (mass%) of organic phase) }

MA partition rate = { MA concentration (mass%) of aqueous phase)/(MA concentration (mass%) of organic phase) }

In the above-mentioned process for producing acetic acid according to the present invention, the liquid separation step satisfies at least one of the following conditions (b-i) to (b-v), and at least a part of the aqueous phase is treated in the acetaldehyde separation and removal step,

(b-ii) the temperature at the time of liquid separation is 70 ℃ or lower;

In the above (b-i), the acetaldehyde concentration in the aqueous phase is, for example, 0.045 to 28.1 mass%, preferably 0.098 to 10 mass%, more preferably 0.098 to 3.0 mass%, particularly preferably 0.098 to 1.0 mass% (e.g., 0.15 to 0.9 mass%). The acetaldehyde concentration in the organic phase is, for example, 0.013 to 24.8 mass%, preferably 0.030 to 2.0 mass%, more preferably 0.030 to 0.50 mass%, particularly preferably 0.030 to 0.24 mass%. In the above (b-i), the acetaldehyde concentration in the aqueous phase is preferably 28.1% by mass or less, and the acetaldehyde concentration in the organic phase is preferably 24.8% by mass or less. In the above (b-i), the acetaldehyde concentration in the liquid supplied in the liquid separation step (for example, the liquid supplied to the decanter 4) is, for example, 26.0 mass% or less (for example, 0.026 to 26.0 mass%), preferably 0.057 to 10 mass%, more preferably 0.057 to 3.0 mass%, particularly preferably 0.057 to 1.0 mass% (for example, 0.057 to 0.42 mass%). In addition, when the acetaldehyde concentration in the aqueous phase exceeds 28.1 mass%, or when the acetaldehyde concentration in the organic phase exceeds 24.8 mass%, the AD distribution ratio becomes small (for example, less than 1.1), and therefore, if the recovery amount of acetaldehyde and the necessity of using an expensive apparatus having extremely high corrosion resistance are compared, the advantage of treating the aqueous phase in the acetaldehyde separation and removal step is extremely small.

In the above (b-ii), the temperature (liquid temperature) at the time of liquid separation is, for example, -5 ℃ to 70 ℃, preferably-5 ℃ to 60 ℃, more preferably-5 ℃ to 51 ℃ (for example, -5 ℃ to 45 ℃), and still more preferably-5 ℃ to 41 ℃ (for example, -5 ℃ to 31 ℃). When the temperature (liquid temperature) during liquid separation exceeds 70 ℃, the AD distribution ratio becomes extremely small, and therefore the advantage of treating the aqueous phase in the acetaldehyde separation and removal step is extremely small.

In the above (b-iii), the methyl acetate concentration in the aqueous phase is, for example, 1.2 to 12.0 mass%, preferably 2.0 to 12.0 mass%, and more preferably 5.0 to 12.0 mass% (e.g., 6.0 to 12.0 mass%). The methyl acetate concentration in the organic phase is, for example, 2.2 to 47.6 mass%, preferably 5.0 to 42 mass%, and more preferably 8.0 to 35 mass% (e.g., 10.0 to 30 mass%). In the above (b-iii), the methyl acetate concentration in the aqueous phase is preferably 12.0 mass% or less and the methyl acetate concentration in the organic phase is preferably 47.6 mass% or less. The sum of the methyl acetate concentration (% by mass) in the aqueous phase and the methyl acetate concentration (% by mass) in the organic phase is, for example, 59.6% by mass or less (for example, 4.2 to 59.6% by mass), preferably 6.0 to 54% by mass, more preferably 8.0 to 54% by mass, still more preferably 10.0 to 54% by mass, particularly preferably 14.0 to 47% by mass (for example, 16.0 to 42% by mass). In the above-mentioned (b-iii), the concentration of methyl acetate in the liquid fed to the liquid separation step (for example, the liquid fed to the decanter 4) is, for example, 38.2% by mass or less (for example, 2.0 to 38.2% by mass), preferably 5.0 to 31% by mass, more preferably 8.0 to 25% by mass, and further preferably 10.0 to 25% by mass). When the methyl acetate concentration in the aqueous phase exceeds 12.0 mass%, when the methyl acetate concentration in the organic phase exceeds 47.6 mass%, or when the sum of the methyl acetate concentration (mass%) in the aqueous phase and the methyl acetate concentration (mass%) in the organic phase exceeds 59.6 mass%, the AD distribution ratio is, for example, less than 1.1, and therefore, for the same reason as described above, the advantage of treating the aqueous phase in the acetaldehyde separation and removal step is extremely small.

In the above (b-iv), the AD distribution ratio is, for example, 1.1 to 8.0, preferably 1.5 to 6.0, and more preferably 1.9 to 5.0. If the AD distribution ratio is less than 1.1, it is extremely industrially disadvantageous to perform the acetaldehyde removal treatment of the aqueous phase, which requires much energy and is liable to corrode the apparatus, because the acetaldehyde concentration in the aqueous phase is low, but if the AD distribution ratio is 1.1 or more (preferably 1.5 or more, more preferably 1.9 or more), the advantage of improving the efficiency of separating and removing acetaldehyde is still large even if an apparatus having high corrosion resistance is used.

In the above (b-v), the MA distribution ratio is 0.25 or more (e.g., 0.25 to 0.70), preferably 0.26 or more (e.g., 0.26 to 0.65), and more preferably 0.28 or more (e.g., 0.28 to 0.60). As described above, the distribution ratio of Methyl Acetate (MA) in the aqueous phase and the organic phase varies depending on the temperature and the composition (components such as acetic acid are contained in addition to water and methyl iodide), and this is also an index for controlling the acetaldehyde distribution ratio.

In the method for producing acetic acid of the present invention, the liquid separation step may satisfy at least one of the conditions (b-i) to (b-v), and 2 or more of the conditions may be satisfied at the same time.

In the process for producing acetic acid according to the present invention, the liquid separation step satisfies at least one of the following conditions (b '-i) to (b' -v), and at least a part of the organic phase is treated in the acetaldehyde separation and removal step,

(b' -ii) the temperature for liquid separation is-5 ℃ or higher;

In the above (b' -i), the acetaldehyde concentration in the aqueous phase is, for example, 0.045 to 35 mass%, preferably 0.15 to 10 mass%, and more preferably 0.2 to 2.0 mass%. The acetaldehyde concentration in the organic phase is, for example, 0.013 to 30% by mass, preferably 0.05 to 5.0% by mass, and more preferably 0.1 to 1.0% by mass. In the above (b' -i), the acetaldehyde concentration in the aqueous phase is preferably 0.045% by mass or more and the acetaldehyde concentration in the organic phase is preferably 0.013% by mass or more. In the above-mentioned (b' -i), the acetaldehyde concentration in the liquid fed to the liquid separation step (for example, the liquid fed to the decanter 4) is, for example, 0.026% by mass or more (for example, 0.026 to 32% by mass), preferably 0.10 to 8.0% by mass, and more preferably 0.15 to 1.8% by mass. When the acetaldehyde concentration in the aqueous phase is less than 0.045% by mass or when the acetaldehyde concentration in the organic phase is less than 0.013% by mass, the AD distribution ratio becomes a large value, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is extremely small.

In the above (b' -ii), the temperature (liquid temperature) at the time of liquid separation is-5 ℃ or higher (e.g., -5 ℃ to 90 ℃), preferably 0 ℃ or higher (e.g., -0 to 90 ℃), more preferably 10 ℃ or higher (e.g., 10 to 90 ℃), further preferably 20 ℃ or higher (e.g., 25 to 90 ℃), a temperature higher than 30 ℃ (e.g., more than 30 ℃ and 90 ℃ or lower)), a temperature higher than 35 ℃ (e.g., more than 35 ℃ and 90 ℃ or lower), a temperature higher than 40 ℃ (e.g., more than 40 ℃ and 90 ℃ or lower), and particularly preferably a temperature higher than 70 ℃ (e.g., more than 70 ℃ and 90 ℃ or lower). When the temperature (liquid temperature) during liquid separation is lower than-5 ℃, the AD distribution ratio is, for example, more than 4.3, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is extremely small.

In the above (b' -iii), the methyl acetate concentration in the aqueous phase is, for example, 1.2 to 20% by mass, preferably 2.5 to 18% by mass, more preferably 4.0 to 15% by mass, still more preferably 6.0 to 13% by mass, and particularly preferably 7.0 to 12% by mass. The methyl acetate concentration in the organic phase is, for example, 2.2 to 60 mass%, preferably 5.8 to 48 mass%, more preferably 8.0 to 40 mass%, still more preferably 10.0 to 30 mass%, and particularly preferably 11.0 to 25 mass%. In the (b' -iii), the methyl acetate concentration in the aqueous phase is preferably 1.2% by mass or more and the methyl acetate concentration in the organic phase is preferably 2.2% by mass or more. The sum of the methyl acetate concentration (% by mass) in the aqueous phase and the methyl acetate concentration (% by mass) in the organic phase is, for example, 3.4 to 75% by mass, preferably 8.3 to 60% by mass (e.g., 10 to 40% by mass), more preferably 15.0 to 50% by mass, and still more preferably 25 to 53% by mass. In the case of (b' -iii), the concentration of methyl acetate in the liquid supplied in the liquid separation step (for example, the liquid supplied to the decanter 4) is, for example, 2.0 to 50% by mass, preferably 5.0 to 38% by mass, more preferably 8.0 to 35% by mass, and still more preferably 10.0 to 32% by mass. When the methyl acetate concentration in the aqueous phase is less than 1.2 mass%, the methyl acetate concentration in the organic phase is less than 2.2 mass%, or the sum of the methyl acetate concentration (mass%) in the aqueous phase and the methyl acetate concentration (mass%) in the organic phase is less than 3.4 mass%, the AD distribution ratio becomes a large value, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is small.

In the above (b' -iv), the AD distribution ratio is 4.1 or less (for example, 0.5 to 4.1), preferably 3.35 or less (for example, 0.6 to 3.35), more preferably 3 or less (0.7 to 3), still more preferably 2.8 or less (for example, 0.8 to 2.8), particularly preferably 2.5 or less (for example, 0.8 to 2.5), particularly preferably 2.3 or less (for example, 0.9 to 2.3), and particularly preferably 2.0 or less (for example, 1.0 to 2.0). When the AD distribution ratio exceeds 4.1, the acetaldehyde concentration in the organic phase is extremely low, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is extremely small. As a method for achieving an AD distribution ratio (less than 1.1) most suitable for the organic phase acetaldehyde removal treatment, for example, there are included: the acetaldehyde concentration in the aqueous phase is made to exceed 28.1 mass%, the acetaldehyde concentration in the organic phase is made to exceed 24.8 mass%, the acetaldehyde concentration in the feed liquid to the liquid separation step is made to exceed 26.0 mass%, the temperature at the time of liquid separation is made to exceed 70 ℃, the methyl acetate concentration in the aqueous phase is made to exceed 12.0 mass%, the methyl acetate concentration in the organic phase is made to exceed 47.6 mass%, the methyl acetate concentration in the feed liquid to the liquid separation step is made to exceed 38.2 mass%, and the like.

In the above (b' -v), the MA partition ratio is 0.8 or less (e.g., 0.15 to 0.80), preferably 0.7 or less (e.g., 0.20 to 0.70), more preferably 0.6 or less (e.g., 0.20 to 0.60), still more preferably 0.44 or less (e.g., 0.20 to 0.44), and particularly preferably less than 0.25 (e.g., 0.20 or more and less than 0.25). The ratio of the Methyl Acetate (MA) in the aqueous phase to the organic phase varies depending on the temperature and composition (including components such as acetic acid in addition to water and methyl iodide), and this also serves as an index for controlling the acetaldehyde distribution ratio.

In the method for producing acetic acid of the present invention, the liquid separation step may satisfy at least one of the conditions (b '-i) to (b' -v), and 2 or more of the above conditions may be satisfied at the same time.

The acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase are determined by the composition of the liquid to be supplied for separation and the temperature at the time of separation. The higher the acetaldehyde concentration in the feed liquid in the liquid separation step, the higher the acetaldehyde concentration in the aqueous phase and the organic phase; the higher the methyl acetate concentration in the feed liquid in the liquid separation step, the higher the methyl acetate concentration in the aqueous phase and the organic phase. As shown in the examples, the higher the temperature at the time of liquid separation, the higher the proportion of acetaldehyde distributed in the organic phase. Further, the concentration of acetaldehyde and the concentration of methyl acetate in the feed liquid in the liquid separation step can be controlled according to, for example, the reaction conditions in the reaction tank 1, the evaporation conditions in the evaporation tank 2, and the distillation conditions in the distillation column 3. In general, the higher the acetaldehyde concentration and the methyl acetate concentration in the reaction mixture, the higher the acetaldehyde concentration and the methyl acetate concentration in the liquid fed to the separation step, respectively. The following tendency was exhibited with respect to the acetaldehyde concentration in the reaction mixture: the reaction temperature, hydrogen partial pressure, methyl iodide concentration, water concentration, catalyst concentration and lithium iodide concentration in the reaction system are increased as they are higher, and the CO partial pressure and methyl acetate concentration are decreased as they are higher, respectively (see jp 2006-182691 a). Further, since methyl acetate is produced by the esterification reaction of acetic acid and methanol, the higher the acetic acid concentration and the methanol concentration in the reaction system are, respectively, the higher the methyl acetate concentration in the reaction mixture is, the more the acetic acid concentration and the methanol concentration are increased, and the higher the water concentration is, the more the water concentration is decreased. In this manner, by adjusting the reaction conditions in the reaction tank and the operating conditions of the evaporation step and the distillation step performed before the liquid separation step, the composition of the liquid supplied in the liquid separation step, and further the acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase can be adjusted.

The separation step may include: an evaporation step of separating the reaction mixture obtained in the carbonylation reaction step into a vapor stream and a residual liquid stream in an evaporation tank; and a low-boiling point removal step of separating the vapor stream into a 1 st overhead stream rich in at least one low-boiling component selected from methyl iodide and acetaldehyde and a 1 st acetic acid stream rich in acetic acid in the 1 st distillation column, and condensing and separating the 1 st overhead stream to obtain an aqueous phase and an organic phase.

The separation step may further include a dehydration step of separating the 1 st acetic acid stream into a 2 nd overhead stream enriched in water and a 2 nd acetic acid stream enriched in acetic acid compared with the 1 st acetic acid stream by a 2 nd distillation column. By dehydrating the 1 st acetic acid stream in the 2 nd distillation column, the 2 nd acetic acid stream having a low water content can be obtained as a column bottom residue or a side stream from the bottom or the middle part of the column. The 2 nd acetic acid stream may be purified directly, or by further purification as desired, to yield product acetic acid.

The reflux ratio of the 2 nd distillation column is, for example, 0.1 or more, preferably 0.3 or more (for example, 0.32 or more), more preferably 1.0 or more, further preferably 5.0 or more, and particularly preferably 10 or more (for example, 12 or more). The upper limit of the reflux ratio of the 2 nd distillation column may be, for example, about 3000 (or 1000) or about 200 (or 100). By increasing the reflux ratio of the 2 nd distillation column to 0.1 or more, the purity of the 2 nd acetic acid stream and the potassium permanganate test value can be improved.

The crotonaldehyde concentration in the 2 nd acetic acid stream is, for example, 0.98 mass ppm or less, preferably 0.80 mass ppm or less, more preferably 0.50 mass ppm or less, and still more preferably 0.30 mass ppm or less. By setting the crotonaldehyde concentration in the 2 nd acetic acid stream to 0.98 mass ppm or less, the crotonaldehyde concentration in the 2 nd acetic acid stream can be significantly reduced, and the potassium permanganate test value of the 2 nd acetic acid stream can be greatly increased. The lower limit of the crotonaldehyde concentration in the 2 nd acetic acid stream may be 0 mass ppm, but may be, for example, 0.01 mass ppm (or 0.10 mass ppm).

The 2-ethylcrotonaldehyde concentration in the 2 nd acetic acid stream is, for example, 1.00 mass ppm or less, preferably 0.50 mass ppm or less, more preferably 0.30 mass ppm or less, and further preferably 0.20 mass ppm or less. The potassium permanganate test value of the 2 nd acetic acid stream can be further improved by adjusting the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream to 1.0 mass ppm or less. The lower limit of the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream may be, for example, 0 mass ppm or 0.01 mass ppm (e.g., 0.10 mass ppm).

The butyl acetate concentration in the 2 nd acetic acid stream is, for example, 15 mass ppm or less, preferably 10 mass ppm or less, more preferably 8 mass ppm or less, and particularly preferably 5 mass ppm or less (for example, 3 mass ppm or less). The purity of the 2 nd acetic acid stream can be improved by setting the butyl acetate concentration in the 2 nd acetic acid stream to 15 mass ppm or less. The lower limit of the butyl acetate concentration in the 2 nd acetic acid stream may be, for example, 0 mass ppm, or 0.1 mass ppm (e.g., 0.3 mass ppm or 1.0 mass ppm).

The potassium permanganate test value of the 2 nd acetic acid stream is preferably more than 50 minutes, more preferably 60 minutes or more, further preferably 100 minutes or more, and particularly preferably 120 minutes or more (for example, 180 minutes or more, particularly preferably 240 minutes or more, and among these, preferably 360 minutes or more). The potassium permanganate test value is also defined in Japanese Industrial Standards (JIS) as an index for controlling the quality of acetic acid as a product. As described above, the potassium permanganate test value is a preferable index for easily confirming the purity of the product acetic acid, which is widely used industrially in quality control of the product acetic acid. The potassium permanganate test can be determined according to the rules of the visual colorimetry of JIS K1351 (1993 version).

The crotonaldehyde concentration in the vapor stream supplied to the 1 st distillation column is, for example, 0 to 5.0 mass ppm (e.g., 0.01 to 4.0 mass ppm), preferably 0.1 to 3.0 mass ppm, and more preferably 0.2 to 2.0 mass ppm. The 2-ethylcrotonaldehyde concentration of the vapor stream is, for example, 0 to 3.0 mass ppm (e.g., 0.01 to 2.5 mass ppm), preferably 0.02 to 2.0 mass ppm, and more preferably 0.03 to 0.8 mass ppm. The butyl acetate concentration of the vapor stream is, for example, 0.1 to 13.0 mass ppm, preferably 0.2 to 12.0 mass ppm, and more preferably 0.3 to 9.0 mass ppm.

The crotonaldehyde concentration in the 1 st acetic acid stream is, for example, 1.3 mass ppm or less, preferably 1.0 mass ppm or less, more preferably 0.85 mass ppm or less, and particularly preferably 0.5 mass ppm or less (for example, 0.25 mass ppm or less). By setting the crotonaldehyde concentration in the 1 st acetic acid stream to 1.3 mass ppm or less, the crotonaldehyde concentration in the 2 nd acetic acid stream can be significantly reduced, and the potassium permanganate test value of the 2 nd acetic acid stream can be significantly increased. The lower limit of the crotonaldehyde concentration in the 1 st acetic acid stream may be 0 mass ppm, but may be, for example, 0.01 mass ppm (or 0.10 mass ppm).

The concentration of 2-ethylcrotonaldehyde in the 1 st acetic acid stream is, for example, 1.0 mass ppm or less, preferably 0.50 mass ppm or less. The potassium permanganate test value of the 2 nd acetic acid stream can be further improved by adjusting the concentration of 2-ethylcrotonaldehyde in the 1 st acetic acid stream to 1.0 mass ppm or less. The lower limit of the 2-ethylcrotonaldehyde concentration in the 1 st acetic acid stream may be, for example, 0 mass ppm, or 0.01 mass ppm (or 0.10 mass ppm).

The butyl acetate concentration in the 1 st acetic acid stream is, for example, 15 mass ppm or less, preferably 10 mass ppm or less, more preferably 8 mass ppm or less, and particularly preferably 5 mass ppm or less (for example, 3 mass ppm or less). By making the butyl acetate concentration in the 1 st acetic acid stream 15 mass ppm or less, the purity of the 2 nd acetic acid stream can be improved. The lower limit of the butyl acetate concentration in the 2 nd acetic acid stream may be, for example, 0 mass ppm, or 0.1 mass ppm (e.g., 0.3 mass ppm or 1.0 mass ppm).

The catalyst system may further contain an ionic iodide. The ionic iodide acts as a promoter.

The above-mentioned method for producing acetic acid has an acetaldehyde separation and removal step. The acetaldehyde separation and removal step is a step of separating and removing acetaldehyde from the process stream, and is a step of selecting and treating an object to be treated from at least a part of the aqueous phase and/or the organic phase based on a liquid temperature at the time of liquid separation in the liquid separation step, an acetaldehyde concentration, an acetaldehyde partition rate, a methyl acetate concentration, or a methyl acetate partition rate of the aqueous phase and/or the organic phase. In a preferred embodiment, the method further comprises an acetaldehyde separation/removal step of distilling at least a part of the aqueous phase and/or the organic phase obtained by condensing the 1 st overhead stream to separate and remove acetaldehyde. In this case, at least a part of the raffinate after the separation of acetaldehyde from at least a part of the aqueous phase and/or the organic phase may be recycled to the reaction tank. By providing the acetaldehyde separation and removal step, acetaldehyde generated in the reaction system can be efficiently separated and removed. Further, by recycling the residual liquid after acetaldehyde separation removal to the reaction tank, useful methyl iodide and the like can be effectively utilized.

The crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is usually 0.01 to 50 mass ppm (e.g., 0.1 to 50 mass ppm), preferably 0.3 to 30 mass ppm, more preferably 0.5 to 10 mass ppm, and still more preferably 0.8 to 7.0 mass ppm (e.g., 1.0 to 5.0 mass ppm).

In the crotonaldehyde removal step, the distillation column is preferably operated so as to satisfy all of the conditions (a-i) to (a-iii) described above. By operating the crotonaldehyde removing column so as to satisfy all of the conditions (a-i) to (a-iii), the crotonaldehyde removal efficiency can be significantly improved, and the potassium permanganate test value of the acetic acid product can be significantly improved.

In the crotonaldehyde removal step, the distillation may be performed by a batch process. If the distillation treatment is carried out batchwise at the time when crotonaldehyde is accumulated to some extent in the aqueous phase and/or the organic phase, energy costs can be saved.

The amount of the distillation column in the crotonaldehyde removing step is, for example, 0.0001 to 50 parts by mass, preferably 0.001 to 30 parts by mass (for example, 0.01 to 10 parts by mass, particularly 0.1 to 5 parts by mass) relative to 100 parts by mass of the amount of the vapor stream supplied to the 1 st distillation column.

The above process for producing acetic acid according to the present invention may further comprise: a washing step of separating an off-gas from the process into a stream enriched in carbon monoxide and a stream enriched in acetic acid by subjecting the off-gas to an absorption treatment with an absorption solvent containing at least acetic acid.

Hereinafter, one embodiment of the present invention will be described. Fig. 1 is an example showing an acetic acid production scheme (methanol carbonylation process) according to an embodiment of the present invention. The acetic acid production apparatus according to this acetic acid production process is provided with a reaction tank 1, an evaporation tank 2, a distillation column 3, a decanter 4, a distillation column 5, a distillation column 6, an ion exchange resin column 7, a washing system 8, an acetaldehyde separation and removal system 9, a distillation column 10,

condensers

1a,2a,3a,5a,6a,10a, a heat exchanger 2b,

reboilers

3b,5b,6b,10b, lines 11 to 56, 58 to 63, and a pump 57, and is configured so as to be able to continuously produce acetic acid. In the method for producing acetic acid according to the present embodiment, the reaction step, the evaporation step (flash evaporation step), the 1 st distillation step, the 2 nd distillation step, the 3 rd distillation step, the crotonaldehyde removal step, and the adsorption removal step are performed in the reaction tank 1, the evaporation tank 2, the distillation column 3, the distillation column 5, the distillation column 6, the distillation column 10, and the ion exchange resin column 7, respectively. The 1 st distillation step is also referred to as a low-boiling point removal step, the 2 nd distillation step is also referred to as a dehydration step, and the 3 rd distillation step is also referred to as a high-boiling point removal step. In addition, the decanter 4 and the acetaldehyde separation and removal system 9 are respectively subjected to a liquid separation step and an acetaldehyde separation and removal (acetaldehyde removal treatment) step. In the present invention, the process is not limited to the above, and for example, the process may be performed without any equipment including the distillation column 5, the distillation column 6, and the ion exchange resin column 7. Further, as will be described later, a product column may be provided downstream of the ion exchange resin column 7.

The reaction tank 1 is a unit for performing a reaction process. This reaction step is a step for continuously producing acetic acid by a reaction (carbonylation reaction of methanol) represented by the following reaction formula (1). In the normal operation state of the acetic acid production apparatus, a reaction mixture stirred by a stirrer, for example, exists in the reaction tank 1. The reaction mixture contains methanol and carbon monoxide as raw materials, a metal catalyst, a promoter, water, acetic acid as a production target, and various by-products, and the liquid phase and the gas phase are in an equilibrium state.

CH ₃ OH+CO→CH ₃ COOH (1)

The raw materials in the reaction mixture were liquid methanol and gaseous carbon monoxide. Methanol is continuously supplied from a methanol storage unit (not shown) to the reaction tank 1 through a line 11 at a predetermined flow rate.

Carbon monoxide is continuously supplied from a carbon monoxide storage unit (not shown) to the reaction tank 1 at a predetermined flow rate through a line 12. The carbon monoxide does not have to be pure carbon monoxide, and may contain other gases such as nitrogen, hydrogen, carbon dioxide, oxygen, etc. in a small amount (for example, 5% by mass or less, preferably 1% by mass or less).

The metal catalyst in the reaction mixture is a substance for promoting the carbonylation of methanol, and for example, a rhodium catalyst, an iridium catalyst may be used. As rhodium catalyst, use may be made of, for example, compounds of the formula [ Rh (CO) ] ₂ I ₂ ] ^- Rhodium complex represented. As iridium catalyst, for example, those represented by the formula [ Ir (CO) ] ₂ I ₂ ] ^- An iridium complex represented. As the metal catalyst, a metal complex catalyst is preferable. The concentration of the catalyst in the reaction mixture (in terms of metal) is, for example, 100 to 10000 mass ppm, preferably 200 to 5000 mass ppm, and more preferably 400 to 3000 mass ppm, relative to the entire liquid phase of the reaction mixture.

The promoter is an iodide for assisting the action of the catalyst, and examples thereof include methyl iodide and ionic iodide. Methyl iodide may exhibit an action of promoting the catalytic action of the above catalyst. The concentration of methyl iodide is, for example, 1 to 20% by mass based on the whole liquid phase of the reaction mixture. The ionic iodide is an iodide (particularly, an ionic metal iodide) which generates an iodide ion in the reaction liquid, and has an effect of stabilizing the catalyst and an effect of suppressing a side reaction. Examples of the ionic iodide include alkali metal iodides such as lithium iodide, sodium iodide, and potassium iodide. The concentration of the ionic iodide in the reaction mixture is, for example, 1 to 25% by mass, preferably 5 to 20% by mass, based on the entire liquid phase of the reaction mixture. In the case of using, for example, an iridium catalyst, a ruthenium compound or an osmium compound may be used as a co-catalyst. The amount of these compounds used is, for example, 0.1 to 30 moles (in terms of metal), preferably 0.5 to 15 moles (in terms of metal) based on 1 mole (in terms of metal) of iridium in total.

Water in the reaction mixture is an essential component for the production of acetic acid in the reaction mechanism of the carbonylation reaction of methanol, and is also an essential component for the solubilization of water-soluble components in the reaction system. The concentration of water in the reaction mixture is, for example, 0.1 to 15% by mass, preferably 0.8 to 10% by mass, more preferably 1 to 6% by mass, and particularly preferably 1.5 to 4% by mass, based on the entire liquid phase of the reaction mixture. The water concentration is preferably 15 mass% or less in order to suppress the energy required for removing water in the acetic acid purification process and promote the efficiency of acetic acid production. In order to control the water concentration, a predetermined flow rate of water may be continuously supplied to the reaction tank 1.

The acetic acid in the reaction mixture comprises: acetic acid fed in advance into the reaction tank 1 before the operation of the acetic acid production apparatus and acetic acid produced as a main product of the carbonylation reaction of methanol. Such acetic acid may function as a solvent in the reaction system. The concentration of acetic acid in the reaction mixture is, for example, 50 to 90% by mass, preferably 60 to 80% by mass, based on the entire liquid phase of the reaction mixture.

As a main by-product contained in the reaction mixture, for example, methyl acetate may be cited. The methyl acetate may be formed by the reaction of acetic acid with methanol. The concentration of methyl acetate in the reaction mixture is, for example, 0.1 to 30% by mass, preferably 1 to 10% by mass, based on the entire liquid phase of the reaction mixture. As a by-product contained in the reaction mixture, hydrogen iodide may also be cited. In view of the reaction mechanism of the carbonylation reaction of methanol, the use of the above-mentioned catalyst and co-catalyst inevitably results in the generation of hydrogen iodide. The concentration of hydrogen iodide in the reaction mixture is, for example, 0.01 to 2 mass% with respect to the entire liquid phase of the reaction mixture.

In addition, as the by-products, for example: hydrogen, methane, carbon dioxide, acetaldehyde, crotonaldehyde, 2-ethylcrotonaldehyde, butyl acetate, dimethyl ether, alkanes, formic acid and propionic acid, and iodoalkanes such as iodoethane, iodopropane, iodobutane, iodohexane and iododecane.

The acetaldehyde concentration in the reaction mixture is, for example, 500 mass ppm or less, preferably 450 mass ppm or less, more preferably 400 mass ppm or less, further preferably 350 mass ppm or less, and particularly preferably 300 mass ppm or less (for example, 250 mass ppm or less). The lower limit of the acetaldehyde concentration in the reaction mixture is, for example, 1 mass ppm (or 10 mass ppm).

The crotonaldehyde concentration in the reaction mixture is, for example, 5 mass ppm or less, preferably 3 mass ppm or less, and more preferably 2 mass ppm or less. The lower limit of the crotonaldehyde concentration in the reaction mixture is 0ppm, but may be 0.1 ppm by mass (or 0.2 ppm by mass), for example. The concentration of 2-ethylcrotonaldehyde in the reaction mixture is, for example, 5 mass ppm or less, preferably 3 mass ppm or less, and more preferably 2 mass ppm or less. The lower limit of the concentration of 2-ethylcrotonaldehyde in the reaction mixture is 0ppm, but may be 0.1 ppm by mass or 0.2 ppm by mass, for example.

In the present invention, as described above, the reflux ratio of the low-boiling-point removal column is controlled to a specific value or more in order to improve the potassium permanganate test value of the product acetic acid. If the reflux ratio of the low-boiling-point removal column is increased, crotonaldehyde can be concentrated to the top of the column. If the concentrated crotonaldehyde is recycled to the reaction tank, the crotonaldehyde is hydrogenated to convert it to butanol, and further, the butanol is reacted with acetic acid to convert it to butyl acetate, which is harmless to the potassium permanganate test. Therefore, in the present invention, the concentration of butyl acetate in the reaction mixture liquid tends to increase. However, an increase in the concentration of butyl acetate sometimes causes a decrease in the purity of acetic acid as a product. Therefore, the concentration of butyl acetate in the reaction mixture is preferably controlled to, for example, 0.1 to 15 mass ppm (particularly 1 to 12 mass ppm, particularly preferably 2 to 9 mass ppm).

In addition, metals [ corrosion metals (also referred to as corrosive metals) ] such as iron, nickel, chromium, manganese, molybdenum, and the like, which are generated due to corrosion of a device, and cobalt, zinc, copper, and the like, which are other metals, may be included in the reaction mixture. The corrosion metal and other metals are also collectively referred to as "corrosion metal and the like".

In the reaction tank 1 in which the above-mentioned reaction mixture is present, the reaction temperature may be set to, for example, 150 to 250 ℃, the reaction pressure on the whole pressure gauge may be set to, for example, 2.0 to 3.5MPa (absolute pressure), and the carbon monoxide partial pressure may be set to, for example, 0.4 to 1.8MPa (absolute pressure), preferably 0.6 to 1.5MPa (absolute pressure).

In a preferred embodiment, the reaction temperature in the reaction vessel 1 in which the above-mentioned reaction mixture is present is set to, for example, 150 to 250 ℃ and the reaction pressure on the whole pressure gauge is set to, for example, 0.4 to 1.8MPa (absolute pressure), preferably 0.6 to 1.6MPa (absolute pressure), and more preferably 0.9 to 1.4MPa (absolute pressure).

The vapor in the gas phase portion in the reaction tank 1 during the operation of the apparatus includes, for example: carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, and the like. The vapor can be extracted from the reaction tank 1 through a line 13. By adjusting the amount of vapor extracted, the pressure in the reaction tank 1 can be controlled, and for example, the pressure in the reaction tank 1 can be kept constant. The vapor extracted from the reaction tank 1 is introduced into a condenser 1a.

In a preferred embodiment, the vapor in the gas phase portion in the reaction tank 1 during the operation of the apparatus includes, for example: carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, and the like. Hydrogen is contained in carbon monoxide used as a raw material In addition, the shift reaction (CO + H) occurs in the reaction tank 1 ₂ O→H ₂ +CO ₂ ) And is generated. The hydrogen partial pressure in the reaction vessel 1 is, for example, 0.001MPa (absolute pressure) or more [ e.g., 0.005MPa or more]Preferably 0.01MPa or more (absolute pressure) [ e.g., 0.015MPa or more]More preferably 0.02MPa (absolute pressure) or more, still more preferably 0.04MPa (absolute pressure) or more, and particularly preferably 0.06MPa (absolute pressure) or more [ e.g., 0.07MPa (absolute pressure) or more]. The upper limit of the hydrogen partial pressure in the reaction vessel is, for example, 0.5MPa (absolute pressure) [ particularly 0.2MPa (absolute pressure) ]]. If the hydrogen partial pressure in the reaction tank is excessively increased, the amount of acetaldehyde produced increases and crotonaldehyde caused by aldol condensation increases, whereas if the hydrogen partial pressure is excessively decreased, the crotonaldehyde → butanol reaction hardly occurs, but the reaction rate greatly fluctuates due to a slight fluctuation of hydrogen, and the operation becomes unstable. The vapor in the gas phase portion in the reaction tank 1 can be extracted from the reaction tank 1 through the line 13. By adjusting the amount of vapor extracted, the pressure in the reaction tank 1 can be controlled, and for example, the pressure in the reaction tank 1 can be kept constant. The vapor extracted from the reaction tank 1 is introduced into a condenser 1a.

The condenser 1a cools and partially condenses the vapor from the reaction tank 1, and separates the vapor into a condensed component and a gas component. The condensed components including, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, etc. are introduced from the condenser 1a into the reaction tank 1 through the line 14 and recycled. Gas components including, for example, carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like are supplied from the condenser 1a to the scrubbing system 8 through the line 15. In the washing system 8, useful components (for example, methyl iodide, water, methyl acetate, acetic acid, etc.) are separated and recovered from the gas components from the condenser 1 a. In the present embodiment, the separation and recovery may be performed by a wet method using an absorption liquid for capturing a useful component in the gas component. The absorption liquid is preferably an absorption solvent containing at least acetic acid and/or methanol. The absorption liquid may contain methyl acetate. For example, as the absorption liquid, a condensed portion of the vapor from the distillation column 6 described later can be used. The separation and recovery may be performed by a pressure swing adsorption method. The separated and recovered useful component (for example, methyl iodide) is introduced from the washing system 8 into the reaction tank 1 through the recycle line 48 and recycled. The gas after the capture of the useful component is discarded via line 49. The gas discharged from the line 49 can be used as a CO source introduced into the bottom of the evaporation tank 2 or the

raffinate recirculation lines

18 and 19 described later. The same applies to the treatment in the washing system 8 and the subsequent recycling and disposal to the reaction tank 1, with respect to the gas components described later supplied from the other condenser to the washing system 8. In the production process of the present invention, it is preferable to have a washing step of separating off-gas from the process into a stream rich in carbon monoxide and a stream rich in acetic acid by subjecting the off-gas to an absorption treatment with an absorption solvent containing at least acetic acid.

Acetic acid is continuously produced as described above in the reaction tank 1 during the operation of the apparatus. The reaction mixture containing acetic acid is continuously withdrawn from the reaction tank 1 at a predetermined flow rate, and then introduced into the subsequent evaporation tank 2 through the line 16.

The evaporation tank 2 is a unit for performing an evaporation step (flash evaporation step). This evaporation step is a step for partially evaporating the reaction mixture continuously introduced into the evaporation vessel 2 through the line 16 (reaction mixture supply line) to separate it into a vapor stream (volatile phase) and a raffinate stream (low-volatile phase). The evaporation may be performed by reducing the pressure of the reaction mixture without heating, or by reducing the pressure of the reaction mixture while heating. In the evaporation step, the temperature of the vapor stream is, for example, 100 to 260 ℃, preferably 120 to 200 ℃, the temperature of the raffinate stream is, for example, 80 to 200 ℃, preferably 100 to 180 ℃, and the pressure in the tank is, for example, 50 to 1000kPa (absolute pressure). The ratio of the vapor stream and the residual liquid stream separated in the evaporation step is, for example, 10/90 to 50/50 (vapor stream/residual liquid stream) in terms of a mass ratio.

The vapor generated in this step includes, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, and the like, and is continuously extracted from the evaporation tank 2 to the line 17 (vapor stream discharge line). A part of the vapor stream withdrawn from the evaporation tank 2 is continuously introduced into the condenser 2a, and the other part of the vapor stream is continuously introduced into the subsequent distillation column 3 through the line 21. The acetic acid concentration of the vapor stream is, for example, 40 to 85 mass% (preferably 50 to 85 mass%), and more preferably 50 to 75 mass% (e.g., 55 to 75 mass%), the methyl iodide concentration is, for example, 2 to 50 mass% (preferably 5 to 30 mass%), the water concentration is, for example, 0.2 to 20 mass% (preferably 1 to 15 mass%), and the methyl acetate concentration is, for example, 0.2 to 50 mass% (preferably 2 to 30 mass%). The residual liquid stream produced in this step contains the catalyst and the co-catalyst (methyl iodide, lithium iodide, etc.) contained in the reaction mixture, water remaining without being volatilized in this step, methyl acetate, acetic acid, formic acid, propionic acid, etc., and is continuously introduced from the evaporation tank 2 to the heat exchanger 2b through the line 18 by using the pump 57. The heat exchanger 2b cools the raffinate stream from the evaporation tank 2. The reduced temperature raffinate stream is continuously introduced from the heat exchanger 2b through the line 19 to the reaction tank 1 and recycled. The

lines

18 and 19 are collectively referred to as a raffinate recycle line. The acetic acid concentration of the raffinate stream is, for example, 55 to 90 mass%, preferably 60 to 85 mass%.

In a preferred embodiment, the vapor generated in the present step contains, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, 2-ethylcrotonaldehyde, butyl acetate, formic acid propionic acid, and iodoalkanes such as iodoethane, iodopropane, iodobutane, iodohexane, and iododecane, and is continuously extracted from the evaporation tank 2 to the line 17 (vapor stream discharge line). A part of the vapor stream withdrawn from the evaporation tank 2 is continuously introduced into the condenser 2a, and the other part of the vapor stream is continuously introduced into the subsequent distillation column 3 through the line 21. The vapor stream has, for example, an acetic acid concentration of 50 to 85 mass%, preferably 55 to 75 mass%, an iodomethane concentration of 2 to 50 mass% (preferably 5 to 30 mass%), a water concentration of 0.2 to 20 mass% (preferably 1 to 15 mass%), and a methyl acetate concentration of 0.2 to 50 mass% (preferably 2 to 30 mass%). The crotonaldehyde concentration of the vapor stream is, for example, 0 to 5.0 mass ppm (e.g., 0.01 to 4.0 mass ppm), preferably 0.1 to 3.0 mass ppm, and more preferably 0.2 to 2.0 mass ppm. The 2-ethylcrotonaldehyde concentration of the vapor stream is, for example, 0 to 3.0 mass ppm (e.g., 0.01 to 2.5 mass ppm), preferably 0.02 to 2.0 mass ppm, and more preferably 0.03 to 0.8 mass ppm. The butyl acetate concentration of the vapor stream is, for example, 0.1 to 13 mass ppm, preferably 0.2 to 12 mass ppm, and more preferably 0.3 to 9 mass ppm.

The raffinate produced in this step contains the catalyst and co-catalyst (methyl iodide, lithium iodide, etc.) contained in the reaction mixture, water remaining without being volatilized in this step, methyl acetate, acetic acid, crotonaldehyde, 2-ethylcrotonaldehyde, butyl acetate, formic acid, propionic acid, etc., and is continuously introduced from the evaporation tank 2 to the heat exchanger 2b through the line 18 by the pump 57. The heat exchanger 2b cools the raffinate stream from the evaporation tank 2. The reduced temperature raffinate stream is continuously introduced from the heat exchanger 2b through the line 19 to the reaction tank 1 and recycled. The

lines

At the bottom of the evaporation tank 2 and/or the raffinate recirculation line (line 18 and/or line 19), a carbon monoxide-containing gas introduction line 54 for introducing a carbon monoxide-containing gas is preferably connected. By introducing carbon monoxide into the raffinate stored in the lower part of the evaporation tank 2 or the raffinate passing through the raffinate recirculation lines 18 and 19 (particularly the line 18), the amount of carbon monoxide dissolved in the raffinate increases, the stability of the catalyst increases, and the settling and accumulation of the catalyst can be prevented. The carbon monoxide content in the introduced carbon monoxide-containing gas is, for example, 10 mass% or more, preferably 20 mass% or more, more preferably 40 mass% or more, and particularly preferably 60 mass% or more.

The condenser 2a partially condenses the vapor stream from the evaporation tank 2 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components include, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, etc., and are introduced from the condenser 2a into the reaction tank 1 through the

lines

22 and 23 and recycled. The gas components, including, for example, carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like, are supplied from the condenser 2a to the scrubbing system 8 through the

lines

20, 15. When the reaction for producing acetic acid in the reaction step is an exothermic reaction, part of the heat accumulated in the reaction mixture is transferred to the vapor generated from the reaction mixture in the evaporation step (flash evaporation step). The condensed component of the vapor produced by cooling in the condenser 2a is recycled to the reaction tank 1. That is, in this acetic acid production apparatus, the heat generated in the carbonylation reaction of methanol can be efficiently removed by the condenser 2 a.

In the preferred embodiment, the condenser 2a partially condenses the vapor stream from the evaporation tank 2 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components include, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, 2-ethyl crotonaldehyde, butyl acetate, formic acid, propionic acid, etc., and are introduced from the condenser 2a into the reaction tank 1 through the

lines

The distillation column 3 is a unit for performing the 1 st distillation step, and may be a so-called low-boiling point removal column in the present embodiment. The 1 st distillation step is a step for separating and removing low-boiling components by subjecting the vapor stream continuously introduced into the distillation column 3 to a distillation treatment. More specifically, in the 1 st distillation step, the vapor stream is distilled and separated into an overhead stream enriched in at least one low-boiling component selected from methyl iodide and acetaldehyde, and an acetic acid stream enriched in acetic acid. The distillation column 3 includes, for example, a plate column and a packed column. When a plate column is used as the distillation column 3, the number of theoretical plates is, for example, 5 to 50 layers, and the reflux ratio is, for example, 0.5 to 3000 depending on the number of theoretical plates. In the distillation column 3, the column top pressure is set to, for example, 80 to 160kPa (gauge pressure), and the column bottom pressure is set to be higher than the column top pressure, for example, 85 to 180kPa (gauge pressure). In the distillation column 3, the column top temperature is set to a temperature of, for example, 90 to 130 ℃ lower than the boiling point of acetic acid at a set column top pressure, and the column bottom temperature is set to, for example, 120 to 160 ℃ higher than the boiling point of acetic acid at a set column bottom pressure.

In a preferred embodiment, the column top pressure is set to, for example, 80 to 160kPa (gauge pressure) and the column bottom pressure is set to be higher than the column top pressure, for example, 85 to 180kPa (gauge pressure) in the distillation column 3. In the distillation column 3, the column top temperature is set to a temperature of, for example, 90 to 130 ℃ lower than the boiling point of acetic acid at the set column top pressure, and the column bottom temperature is set to, for example, 115 to 165 ℃ (preferably 120 to 160 ℃) higher than the boiling point of acetic acid at the set column bottom pressure.

The vapor stream from the evaporation tank 2 is continuously introduced into the distillation column 3 through the line 21, and the vapor as an overhead stream is continuously extracted from the top of the distillation column 3 to the line 24. From the bottom of the distillation column 3, a still bottoms is continuously withdrawn to line 25. And 3b is a reboiler. From a height position between the top and bottom of the distillation column 3, an acetic acid stream (1 st acetic acid stream; liquid) as a side stream is continuously withdrawn from the line 27.

The vapor withdrawn from the column top of the distillation column 3 contains more components having a boiling point lower than that of acetic acid (low-boiling components) than the above-mentioned bottom residue and side stream from the distillation column 3, including, for example: methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, acetaldehyde, formic acid, and the like. The vapor also contains acetic acid. Such vapor is continuously introduced into the condenser 3a through a line 24.

In a preferred embodiment, components having a boiling point lower than that of acetic acid (low boiling components) are contained more in the vapor withdrawn from the column top of the distillation column 3 than the above-mentioned still bottoms and side streams from the distillation column 3, including, for example: methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, formic acid and the like. The vapor also contains acetic acid. Such vapor is continuously introduced into the condenser 3a through the line 24.

The condenser 3a partially condenses the vapor from the distillation column 3 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components, including, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, etc., are continuously introduced from the condenser 3a into the decanter 4 through the line 28. The condensed components introduced into the decanter 4 are separated into an aqueous phase (upper phase) and an organic phase (methyl iodide phase; lower phase). The aqueous phase contains water, and for example, methyl iodide, hydrogen iodide, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like. The organic phase contains methyl iodide, and for example, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like. In the present embodiment, a part of the aqueous phase is refluxed to the distillation column 3 through the line 29, and the other part of the aqueous phase is introduced to the reaction tank 1 through the

lines

29, 30, and 23 and recycled. A part of the organic phase is introduced into the reaction tank 1 through the lines 31, 23 and recycled. The other part of the organic phase and/or the other part of the aqueous phase is introduced into the acetaldehyde separation and removal system 9 through the lines 31,50 and/or the

lines

30, 51.

In a preferred embodiment, the condenser 3a partially condenses the vapor from the distillation column 3 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components including, for example, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, and formic acid are continuously introduced from the condenser 3a into the decanter 4 through the line 28. The condensed components introduced into the decanter 4 are separated into an aqueous phase (upper phase) and an organic phase (methyl iodide phase; lower phase). The aqueous phase contains water, and for example, methyl iodide, hydrogen iodide, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, formic acid, and the like. The organic phase contains methyl iodide, and for example, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, and formic acid.

In the present invention, the phase to be supplied to the acetaldehyde separation and removal step can be determined based on the acetaldehyde concentration in the aqueous phase, the acetaldehyde concentration in the organic phase, the methyl acetate concentration in the aqueous phase, the methyl acetate concentration in the organic phase, the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase, the temperature (liquid temperature) at the time of liquid separation, the acetaldehyde distribution ratio (AD distribution ratio) obtained by the following equation, or the methyl acetate distribution ratio (MA distribution ratio) obtained by the following equation in a liquid separation step (for example, liquid separation by a decanter 4) of separating a process stream containing at least water, acetic Acid (AC), methyl iodide (MeI), and Acetaldehyde (AD) into an aqueous phase and an organic phase. From the other side, suitable conditions for separation when the organic phase is supplied to the acetaldehyde separation and removal step and suitable conditions for separation when the aqueous phase is supplied to the acetaldehyde separation and removal step are shown.

(b-ii) the temperature for liquid separation is 70 ℃ or lower;

(b-iv) an acetaldehyde partition rate [ { acetaldehyde concentration (mass%) of the aqueous phase }/{ acetaldehyde concentration (mass%) of the organic phase } is 1.1 or more;

(b-v) a methyl acetate partition rate [ { methyl acetate concentration (mass%) of the aqueous phase }/{ methyl acetate concentration (mass%) of the organic phase } ] of 0.25 or more.

In the above (b-iii), the methyl acetate concentration in the aqueous phase is, for example, 1.2 to 12.0 mass%, preferably 2.0 to 12.0 mass%, and more preferably 5.0 to 12.0 mass% (e.g., 6.0 to 12.0 mass%). The methyl acetate concentration in the organic phase is, for example, 2.2 to 47.6 mass%, preferably 5.0 to 42 mass%, and more preferably 8.0 to 35 mass% (e.g., 10.0 to 30 mass%). In the above (b-iii), the methyl acetate concentration in the aqueous phase is preferably 12.0% by mass or less and the methyl acetate concentration in the organic phase is preferably 47.6% by mass or less. The sum of the methyl acetate concentration (% by mass) in the aqueous phase and the methyl acetate concentration (% by mass) in the organic phase is, for example, 59.6% by mass or less (for example, 4.2 to 59.6% by mass), preferably 6.0 to 54% by mass, more preferably 8.0 to 54% by mass, still more preferably 10.0 to 54% by mass, particularly preferably 14.0 to 47% by mass (for example, 16.0 to 42% by mass). In the above-mentioned (b-iii), the concentration of methyl acetate in the liquid fed to the liquid separation step (for example, the liquid fed to the decanter 4) is, for example, 38.2% by mass or less (for example, 2.0 to 38.2% by mass), preferably 5.0 to 31% by mass, more preferably 8.0 to 25% by mass, and further preferably 10.0 to 25% by mass). When the methyl acetate concentration in the aqueous phase exceeds 12.0 mass%, when the methyl acetate concentration in the organic phase exceeds 47.6 mass%, or when the sum of the methyl acetate concentration (mass%) in the aqueous phase and the methyl acetate concentration (mass%) in the organic phase exceeds 59.6 mass%, the AD distribution ratio is, for example, less than 1.1, and therefore, for the same reason as described above, the advantage of treating the aqueous phase in the acetaldehyde separation and removal step is extremely small.

In the above (b-v), the MA distribution ratio is 0.25 or more (e.g., 0.25 to 0.70), preferably 0.26 or more (e.g., 0.26 to 0.65), and more preferably 0.28 or more (e.g., 0.28 to 0.60). As described above, the distribution ratio of Methyl Acetate (MA) in the aqueous phase and the organic phase varies depending on the temperature and composition (including components such as acetic acid in addition to water and methyl iodide), and this also serves as an index for controlling the acetaldehyde distribution ratio.

In the method for producing acetic acid of the present invention, the liquid separation step may satisfy at least one of the conditions (b-i) to (b-v), but 2 or more of the above conditions may be satisfied at the same time. As a combination of 2 or more conditions preferably satisfied simultaneously, the following combinations can be cited: (b-i) and (b-ii), (b-i) and (b-iii), (b-i) and (b-iv), (b-i) and (b-v), (b-ii) and (b-iii), (b-ii) and (b-iv), (b-ii) and (b-v), (b-iii) and (b-iv), (b-iii) and (b-v), (b-iv) and (b-v), (b-i) and (b-ii) and (b-iii), (b-i) and (b-ii) and (b-iv), (b-i) and (b-ii) and (b-v), (b-i) and (b-iii) and (b-iv), (b-i) and (b-iii) and (b-v), (b-i) and (b-iv) and (b-v), (b-ii) and (b-iii) and (b-iv), (b-ii) and (b-iv) and (b-v), (b-ii) and (b-iv) and (b-v), (b-iii) and (b-iv) and (b-v), (b-i) and (b-ii) and (b-iii) and (b-iv), (b-i) and (b-ii) and (b-iii) and (b-v), (b-i) and (b-iii) and (b-iv) and (b-v), (b-ii) and (b-iii) and (b-iv) and (b-v), and combinations of (b-i) and (b-ii) and (b-iii) and (b-iv) and (b-v). Among them, it is particularly preferable that at least (b-i), (b-ii) and (b-iii) are satisfied at the same time, that at least (b-i), (b-ii), (b-iii) and (b-iv) are satisfied at the same time, or that all of (b-i), (b-ii), (b-iii), (b-iv) and (b-v) are satisfied at the same time.

On the other hand, in the above-mentioned method for producing acetic acid according to the present invention, the liquid separation step satisfies at least one of the following conditions (b '-i) to (b' -v), and at least a part of the organic phase is treated in the acetaldehyde separation and removal step,

(b' -ii) the temperature for liquid separation is-5 ℃ or higher;

In the above (b' -iii), the methyl acetate concentration in the aqueous phase is, for example, 1.2 to 20 mass%, preferably 2.5 to 18 mass%, more preferably 4.0 to 15 mass%, further preferably 6.0 to 13 mass%, particularly preferably 7.0 to 12 mass%. The methyl acetate concentration in the organic phase is, for example, 2.2 to 60 mass%, preferably 5.8 to 48 mass%, more preferably 8.0 to 40 mass%, still more preferably 10.0 to 30 mass%, particularly preferably 11.0 to 25 mass%. In the (b' -iii), the methyl acetate concentration in the aqueous phase is preferably 1.2% by mass or more and the methyl acetate concentration in the organic phase is preferably 2.2% by mass or more. The sum of the methyl acetate concentration (% by mass) in the aqueous phase and the methyl acetate concentration (% by mass) in the organic phase is, for example, 3.4 to 75% by mass, preferably 8.3 to 60% by mass (e.g., 10 to 40% by mass), more preferably 15.0 to 50% by mass, and still more preferably 25 to 53% by mass. In the case of (b' -iii), the concentration of methyl acetate in the liquid supplied in the liquid separation step (for example, the liquid supplied to the decanter 4) is, for example, 2.0 to 50% by mass, preferably 5.0 to 38% by mass, more preferably 8.0 to 35% by mass, further preferably 10.0 to 32% by mass, and particularly preferably 15.0 to 31% by mass. When the methyl acetate concentration in the aqueous phase is less than 1.2 mass%, the methyl acetate concentration in the organic phase is less than 2.2 mass%, or the sum of the methyl acetate concentration (mass%) in the aqueous phase and the methyl acetate concentration (mass%) in the organic phase is less than 3.4 mass%, the AD distribution ratio becomes a large value, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is small.

In the above (b' -iv), the AD distribution ratio is 4.1 or less (e.g., 0.5 to 4.1), preferably 3.35 or less (e.g., 0.6 to 3.35), more preferably 3 or less (0.7 to 3), further preferably 2.8 or less (e.g., 0.8 to 2.8), particularly preferably 2.5 or less (e.g., 0.8 to 2.5), particularly preferably 2.3 or less (e.g., 0.9 to 2.3), and particularly preferably 2.0 or less (e.g., 1.0 to 2.0). When the AD distribution ratio exceeds 4.1, the acetaldehyde concentration in the organic phase is extremely low, and therefore, the advantage of treating the organic phase in the acetaldehyde separation and removal step is extremely small. As a method for achieving an AD distribution ratio (less than 1.1) most suitable for the organic phase acetaldehyde removal treatment, for example, there are included: the acetaldehyde concentration in the aqueous phase is made to exceed 28.1 mass%, the acetaldehyde concentration in the organic phase is made to exceed 24.8 mass%, the acetaldehyde concentration in the liquid feed to the liquid separation step is made to exceed 26.0 mass%, the temperature at the time of liquid separation is made to exceed 70 ℃, the methyl acetate concentration in the aqueous phase is made to exceed 12.0 mass%, the methyl acetate concentration in the organic phase is made to exceed 47.6 mass%, the methyl acetate concentration in the liquid feed to the liquid separation step is made to exceed 38.2 mass%, and the like.

In the method for producing acetic acid of the present invention, the liquid separation step may satisfy at least one of the conditions (b '-i) to (b' -v), and 2 or more of the above conditions may be satisfied at the same time. As a combination of 2 or more conditions preferably satisfied simultaneously, the following combinations can be cited: (b '-i) and (b' -ii), (b '-i) and (b' -iii), (b '-i) and (b' -iv), (b '-i) and (b' -v), (b '-ii) and (b' -iii), (b '-ii) and (b' -iv), (b '-ii) and (b' -v), (b '-iii) and (b' -iv), (b '-iii) and (b' -v), (b '-iv) and (b' -v), (b '-i) and (b' -ii) and (b '-iii), (b' -i) and (b '-ii) and (b' -iv), (b '-i) and (b' -ii) and (b '-v), (b' -i) and (b '-iii) and (b' -iv), (b '-i) and (b' -iii) and (b '-v), (b' -i) and (b '-iv) and (b' -v), (b '-ii) and (b' -iv), (b '-ii) and (b' -iii) and (b '-v), (b' -ii) and (b '-iv) and (b' -v), (b '-iii) and (b' -iv) and (b '-v), (b' -i) and (b '-ii) and (b' -iii) and (b '-iv), (b' -i) and (b '-ii) and (b' -iii) and (b '-v), (b' -i) and (b '-iii) and (b' -iv) and (b '-v), (b' -ii) and (b '-iii) and (b' -iv) and (b '-v), and a combination of (b' -i) and (b '-ii) and (b' -iii) and (b '-iv) and (b' -v). Among them, it is particularly preferable that at least (b '-i), (b' -ii) and (b '-iii) are satisfied at the same time, at least (b' -i), (b '-ii), (b' -iii) and (b '-iv) are satisfied at the same time, or all of (b' -i), (b '-ii), (b' -iii), (b '-iv) and (b' -v) are satisfied at the same time.

The acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase are determined by the composition of the liquid supplied in the liquid separation step and the temperature at the time of liquid separation. The higher the acetaldehyde concentration in the feed liquid in the liquid separation step, the higher the acetaldehyde concentration in the aqueous phase and the organic phase; the higher the methyl acetate concentration in the feed liquid in the liquid separation step, the higher the methyl acetate concentration in the aqueous phase and the organic phase. As shown in the examples, the higher the temperature at the time of liquid separation, the higher the proportion of acetaldehyde distributed in the organic phase. Further, the concentration of acetaldehyde and the concentration of methyl acetate in the feed liquid in the liquid separation step can be controlled according to, for example, the reaction conditions in the reaction tank 1, the evaporation conditions in the evaporation tank 2, and the distillation conditions in the distillation column 3. In general, the higher the acetaldehyde concentration and the methyl acetate concentration in the reaction mixture, the higher the acetaldehyde concentration and the methyl acetate concentration in the liquid fed to the separation step, respectively. The following tendency was exhibited with respect to the acetaldehyde concentration in the reaction mixture: the reaction temperature, the hydrogen partial pressure, the methyl iodide concentration, the water concentration, the catalyst concentration, and the lithium iodide concentration in the reaction system are increased as they are higher, and the CO partial pressure and the methyl acetate concentration are decreased as they are higher, respectively (see japanese patent laid-open No. 2006-182691). Further, since methyl acetate is produced by the esterification reaction of acetic acid and methanol, the higher the acetic acid concentration and the methanol concentration in the reaction system are, respectively, the higher the methyl acetate concentration in the reaction mixture is, the more the acetic acid concentration and the methanol concentration are increased, and the higher the water concentration is, the more the water concentration is decreased. In this manner, by adjusting the reaction conditions in the reaction tank and the operating conditions of the evaporation step and the distillation step performed before the liquid separation step, the composition of the liquid supplied in the liquid separation step, and further the acetaldehyde concentration and the methyl acetate concentration in the aqueous phase and the organic phase can be adjusted.

Further, the methyl acetate concentration of the feed to the low-boiling point removal column can be controlled by controlling the methyl acetate concentration in the reaction tank. For reaction tankThe methyl acetate concentration of (2) can be determined by, for example, floating the temperature, rhodium complex (Rh), methyl iodide (MeI), and H in the reaction tank up and down under the condition that the feed amounts of methanol and CO are constant ₂ O, hydrogen partial pressure, carbon monoxide partial pressure, lithium iodide (LiI), and other factors that affect the reaction rate. I.e. if increasing the temperature, rh, meI, H ₂ O, hydrogen partial pressure, carbon monoxide partial pressure, liI, etc., the methyl acetate concentration in the reaction tank decreases, and if it decreases, the methyl acetate concentration increases. In addition, if these conditions are made constant, namely, the temperature, rh, meI, H ₂ Under the condition of constant O, hydrogen partial pressure, carbon monoxide partial pressure, liI and the like, when the feeding amount of methanol and CO is increased, the concentration of methyl acetate is increased, and when the feeding amount of methanol and CO is decreased, the concentration of methyl acetate is decreased. Since the operation is performed under a certain condition with respect to the methyl acetate concentration in the vapor (feed to the low-boiling point removal column) obtained by evaporating the reaction liquid in the evaporation tank, the evaporation rate changes in proportion to the methyl acetate concentration in the reaction liquid. When the evaporation tank is heated and cooled, the methyl acetate concentration also changes due to a change in the evaporation rate. For example, when heating is performed, the acetic acid concentration in the vapor in the evaporation tank increases and the methyl acetate concentration decreases. On the other hand, when cooling is performed, the opposite phenomenon occurs, that is, the acetic acid concentration in the vapor in the evaporation tank decreases and the methyl acetate concentration increases.

In the acetaldehyde separation and removal step using the acetaldehyde separation and removal system 9, the acetaldehyde is separated and removed from the process stream containing at least water, acetic Acid (AC), methyl iodide (MeI), and Acetaldehyde (AD). As a method for separating and removing acetaldehyde, a known method can be used. For example, acetaldehyde is separated off by distillation, extraction or a combination of these. The separated acetaldehyde is discharged to the outside of the apparatus through a line 53. In addition, useful components (e.g., methyl iodide, etc.) contained in the organic phase and/or the aqueous phase are recycled to the reaction tank 1 through the lines 52, 23 to be reused.

In the present embodiment, a part of the aqueous phase is refluxed to the distillation column 3 through the line 29, and the other part of the aqueous phase is introduced into the reaction tank 1 through the

lines

29, 30, and 23 and recycled. A part of the organic phase is introduced into the reaction tank 1 through the lines 31, 23 and recycled. The other part of the organic phase and/or the other part of the aqueous phase are introduced into the acetaldehyde separation and removal system 9 via lines 31, 50, and/or

lines

30, 51. In addition to or instead of refluxing the aqueous phase, a part of the organic phase may be refluxed to the distillation column 3.

The reflux ratio of the distillation column 3 is explained below. When only the aqueous phase of the condensed component of the overhead stream (1 st overhead stream) is refluxed in the distillation column 3, the reflux ratio of the aqueous phase (reflux amount of the aqueous phase/distilled amount of the aqueous phase) is preferably 2 or more, preferably 3 or more, more preferably 4 or more, further preferably 8 or more, and particularly preferably 10 or more. In the case where only the organic phase of the condensed component in the overhead stream is refluxed in the distillation column 3, the reflux ratio of the organic phase (the amount of reflux of the organic phase/the amount of distillate of the organic phase) is preferably 1 or more, preferably 1.5 or more, more preferably 2 or more, further preferably 4 or more, and particularly preferably 5 or more. When both the aqueous phase and the organic phase of the condensed component in the overhead stream in the distillation column 3 are refluxed, the total reflux ratio of the aqueous phase and the organic phase (the total reflux amount of the aqueous phase and the organic phase/the total distillate amount of the aqueous phase and the organic phase) is preferably, for example, 1.5 or more, preferably 2.3 or more, more preferably 3 or more, further preferably 6 or more, and particularly preferably 7.5 or more. When the aqueous phase is refluxed in the distillation column 3, the reflux ratio of the aqueous phase (reflux amount of the aqueous phase/distilled amount of the aqueous phase) is preferably 2 or more, more preferably 3 or more, further preferably 5 or more, particularly preferably 8 or more, and particularly preferably 12 or more. The upper limit of the reflux ratio of the distillation column 3 may be, for example, 3000 (particularly 1000) or 100 (particularly 30) in any case. Since crotonaldehyde (boiling point 104 ℃) has a boiling point lower than that of acetic acid (boiling point 117 ℃), crotonaldehyde can be further concentrated to the top of the distillation column 3 by increasing the reflux ratio of the distillation column 3, and thus the concentration of crotonaldehyde in the 1 st acetic acid stream obtained as a side stream, for example, is reduced. When the condensed components (aqueous phase and/or organic phase) of the 1 st overhead stream in which crotonaldehyde is concentrated are recycled to the reaction tank 1 by increasing the reflux ratio of the distillation column 3, crotonaldehyde reacts with acetaldehyde in the reaction tank 1 to produce 2-ethyl crotonaldehyde. Further, crotonaldehyde reacts with hydrogen in the reaction tank 1 to produce butanol, and this butanol reacts with acetic acid to produce butyl acetate. Compared with crotonaldehyde, the 2-ethyl crotonaldehyde has smaller influence on a potassium permanganate test value, and butyl acetate cannot influence the potassium permanganate test value completely. Therefore, the quality of acetic acid tends to be further improved. Since the boiling points of 2-ethylcrotonaldehyde and butyl acetate are 137 ℃ and 126 ℃ respectively, which are higher than the boiling point of acetic acid (117 ℃), if the reflux ratio of the distillation column 3 is increased, the overhead concentration decreases, and the overhead and bottom liquid are easily concentrated to a side stream located above the feed liquid supply position of the distillation column 3.

Further, in the present embodiment, a part of the organic phase is introduced into the distillation column 10 (crotonaldehyde removing column) through the

lines

31, 50, and 58, and crotonaldehyde is separated and removed by distillation. The distillation may be carried out by any of a continuous type (continuous operation) and a batch type (batch treatment). When the amount of crotonaldehyde produced in the reaction system is very small, it is preferable to separate and remove crotonaldehyde by a batch process at the time when crotonaldehyde is accumulated to some extent in the aqueous phase or the organic phase, for energy cost saving and the like. In addition, when the continuous operation is performed, the quality can be maintained and the steam can be saved by changing the treatment amount (feed amount). The amount of the solvent to be treated in the distillation column 10 (crotonaldehyde removal column) may be, for example, 0.0001 to 50 parts by mass (for example, 0.001 to 30 parts by mass) or 0.01 to 10 parts by mass (for example, 0.1 to 5 parts by mass) when the amount of the solvent fed into the distillation column 3 (the 1 st distillation column; the low-boiling point removal column) is 100 parts by mass. The distillation column 10 includes a rectifying column such as a plate column and a packed column. The theoretical plates of the distillation column 10 are, for example, 1 to 100 layers, preferably 2 to 50 layers, more preferably 4 to 30 layers, and still more preferably 5 to 20 layers (for example, 6 to 15 layers). In the case of continuous distillation, the feed position of the feed liquid to the distillation column 10 is preferably at a position intermediate in the height direction of the distillation column (between the 1 st layer below the column top and the 1 st layer above the column bottom), but may be about 20% to 80% (2/10 to 8/10) downward from the top. If the feed position is too low, the loss of methyl iodide increases, and if the feed position is too high, the crotonaldehyde removal amount (and the alkane removal amount) decreases. The crotonaldehyde concentration in the feed liquid (feed liquid) to the distillation column 10 is usually 0.01 to 50 mass ppm (e.g., 0.1 to 50 mass ppm), preferably 0.3 to 30 mass ppm, more preferably 0.5 to 10 mass ppm, and still more preferably 0.8 to 7.0 mass ppm (e.g., 1.0 to 5.0 mass ppm). The overhead vapor of the distillation column 10 is introduced into the condenser 10a via the line 59 and condensed. A part of the condensate is refluxed to the distillation column 10 via a line 61, and the remaining amount of the condensate is extracted as a distillate via a line 62. The distillate mainly comprises methyl iodide and methyl acetate, and also comprises dimethyl ether, low-boiling alkanes and the like. The distillate can be recycled to the decanter 4 or the reaction tank 1, for example. The gaseous components of the overhead vapor not condensed by condenser 10a are sent via line 63, for example, to scrubbing system 8. From the bottom of the distillation column 10, a still bottoms can be extracted via line 60. The residual liquid in the tower bottom mainly contains high boiling point impurities such as crotonaldehyde, 2-ethyl crotonaldehyde, butyl acetate, alkanes and the like, and acetic acid. The bottoms are usually discarded. The water contained in the organic phase in a trace amount can be concentrated to the top of the column or extracted from the bottom of the column. In addition to the organic phase being introduced into the distillation column 10, or instead of introducing the organic phase into the distillation column 10, the aqueous phase may be introduced into the distillation column 10 through the

lines

30,51, 58. In this case, a distillate containing water is obtained from the top of the distillation column 10, and a bottom residue containing high boiling impurities such as crotonaldehyde and acetic acid is obtained from the bottom. By treating the aqueous phase and/or the organic phase in the distillation column 10 in this manner, crotonaldehyde can be efficiently removed, and the potassium permanganate test value of the product acetic acid can be improved, so that large-scale facilities such as ozone treatment facilities can be eliminated, or small-scale facilities, steam costs, and electricity costs can be reduced. The reflux ratio (reflux amount/distillate amount) of the distillation column 10 is, for example, 0.01 or more, preferably 0.05 or more, more preferably 0.5 or more, further preferably 5 or more, and particularly preferably 20 or more (for example, 30 or more). The upper limit of the reflux ratio of the distillation column 10 is, for example, 1000 (or 100). When the reflux ratio of the distillation column 10 is too large, crotonaldehyde which has been condensed at the bottom of the column is condensed to the top of the column, and the concentration of acetic acid having a higher boiling point increases, so that the reflux ratio of the distillation column 10 is preferably 100 or less. Since crotonaldehyde is extracted from the bottom of the column, the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate of the distillation column 10 to the crotonaldehyde concentration (mass ppm) in the feed liquid is, for example, less than 1, preferably 0.95 or less, more preferably 0.80 or less, still more preferably 0.70 or less, and particularly preferably 0.60 or less (for example, 0.50 or less, particularly 0.30 or less, and among them 0.20 or less). The ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the bottom residual liquid of the distillation column 10 to the crotonaldehyde concentration (mass ppm) in the feed liquid is, for example, greater than 1, preferably 1.2 or more, more preferably 1.5 or more, still more preferably 2.0 or more, particularly preferably 3.0 or more (e.g., 4.0 or more, particularly 5.0 or more), and particularly 10 or more (e.g., 20 or more).

In the preferred embodiment, in the acetaldehyde separation and removal step using the acetaldehyde separation and removal system 9, acetaldehyde contained in the organic phase and/or the aqueous phase is separated and removed by a known method, for example, distillation, extraction, or a combination thereof. The separated acetaldehyde is discharged to the outside of the apparatus through a line 53. In addition, useful components (e.g., methyl iodide, etc.) contained in the organic phase and/or the aqueous phase are recycled to the reaction tank 1 through the lines 52, 23 to be reused.

FIG. 2 is a schematic flow chart showing an example of the system for separating and removing acetaldehyde. According to this flow scheme, for example, when the organic phase is treated in the acetaldehyde separation and removal step, the organic phase is fed to a distillation column (1 st acetaldehyde removal column) 91 via a line 101 and distilled to be separated into an acetaldehyde-enriched overhead stream (line 102) and a methyl iodide-enriched raffinate stream (line 103). The overhead stream is condensed by a condenser 91a, a portion of the condensate is refluxed to the top of the distillation column 91 (line 104), and the other portion of the condensate is fed to the extraction column 92 (line 105). The condensate supplied to the extraction column 92 is subjected to an extraction treatment with water introduced via line 109. The extract obtained by the extraction treatment is fed through a line 107 to a distillation column (2 nd acetaldehyde removal column) 93 and distilled to be separated into an acetaldehyde-enriched overhead stream (line 112) and a water-enriched raffinate stream (line 113). The acetaldehyde-enriched overhead stream is then condensed by condenser 93a, and a portion of the condensate is refluxed to the top of distillation column 93 (line 114) and the other portion of the condensate is discharged outside the system (line 115). In addition, the methyl iodide-rich raffinate stream as bottoms from the 1 st de-acetalization column 91, the methyl iodide-rich raffinate stream obtained in the extraction column 92 (line 108), and the water-rich raffinate stream as bottoms from the 2 nd de-acetalization column 93 are each recycled to the reaction tank 1 via

lines

103, 111, 113, or to a suitable part of the process for reuse. For example, the methyl iodide-enriched raffinate from the extraction column 92 may be recycled to the distillation column 91 via line 110. The liquid in 113 is usually discharged as drainage to the outside. The gas that is not condensed in the

condensers

91a and 93a (lines 106 and 116) is subjected to an absorption treatment or a disposal treatment in the washing system 8.

In the case where the aqueous phase is treated in the acetaldehyde separation and removal step according to the flow of fig. 2, for example, the aqueous phase is fed to a distillation column (1 st acetaldehyde removal column) 91 through a line 101 and distilled to be separated into an acetaldehyde-rich overhead stream (line 102) and a water-rich raffinate stream (line 103). The overhead stream is condensed by condenser 91a, a portion of the condensate is refluxed to the top of distillation column 91 (line 104), and the other portion of the condensate is fed to extraction column 92 (line 105). The condensate supplied to the extraction column 92 is subjected to an extraction treatment using water introduced through line 109. The extract obtained by the extraction treatment is supplied to a distillation column (2 nd acetaldehyde-removing column) 93 via a line 107 to be distilled and separated into an acetaldehyde-rich overhead stream (line 112) and a water-rich raffinate stream (line 113). The acetaldehyde-enriched overhead stream is then condensed by condenser 93a, and a portion of the condensate is refluxed to the top of distillation column 93 (line 114) and the other portion of the condensate is discharged outside the system (line 115). In addition, the water-enriched raffinate stream as bottoms from the 1 st de-acetalization column 91, the methyl iodide-enriched raffinate stream obtained in the extraction column 92 (line 108), and the water-enriched raffinate stream as bottoms from the 2 nd de-acetalization column 93 are each recycled to the reaction tank 1 via

lines

103, 111, 113, or to a suitable part of the process for reuse. For example, the methyl iodide-enriched raffinate from the extraction column 92 may be recycled to the distillation column 91 via line 110. The liquid in 113 is usually discharged to the outside as drainage. The gas that is not condensed by the

condensers

91a and 93a (lines 106 and 116) is subjected to an absorption treatment in the scrubbing system 8 or is subjected to a disposal treatment.

As for the above acetaldehyde from the process stream containing at least water, acetic Acid (AC), methyl iodide (MeI) and Acetaldehyde (AD), it can be separated off by extractive distillation in addition to the above-mentioned method. For example, the organic phase and/or the aqueous phase (feed liquid) obtained by separating the process stream may be supplied to a distillation column (extractive distillation column), an extraction solvent (usually water) may be introduced into a concentration region in the distillation column in which methyl iodide and acetaldehyde are concentrated (for example, a space from the top of the column to a feed liquid supply position), a liquid (extract) descending from the concentration region may be extracted as a side stream (side stream), the side stream may be separated into an aqueous phase and an organic phase, and the aqueous phase may be distilled to discharge acetaldehyde to the outside of the system. In the case where a large amount of water is present in the distillation column, the liquid descending from the concentration region may be extracted as a side stream without introducing the extraction solvent into the distillation column. For example, a unit (chimney tray or the like) capable of receiving the liquid (extract) descending from the concentration region may be disposed in the distillation column, and the liquid (extract) received in the unit may be extracted as a side stream. The introduction position of the extraction solvent is preferably above the feeding position of the feed liquid, and more preferably near the top of the column. The position of the side stream extraction is preferably below the position of the extraction solvent and above the position of the feed liquid supply in the height direction of the column. According to this method, acetaldehyde can be extracted at a high concentration from a concentrate of methyl iodide and acetaldehyde by using an extraction solvent (usually water), and acetaldehyde can be efficiently extracted with a small amount of the extraction solvent because the extraction solvent can be used as an extraction region between an introduction portion and a side stream portion of the extraction solvent. Therefore, for example, compared with a method of extracting an extract liquid by extractive distillation from the bottom of a distillation column (extractive distillation column), the number of layers of the distillation column can be greatly reduced, and the vapor load can be reduced. Further, since a small amount of the extraction solvent is used, the ratio of methyl iodide to acetaldehyde (MeI/AD ratio) in the aqueous extract can be reduced as compared with the method in which the above-described acetaldehyde removal distillation and aqueous extraction of fig. 2 are combined, and therefore acetaldehyde can be removed under the condition that the loss of methyl iodide to the outside of the system is suppressed. The acetaldehyde concentration in the side stream is significantly higher than the acetaldehyde concentration in the feed and the bottoms (bottoms). Further, the ratio of acetaldehyde to methyl iodide in the side stream is larger than the ratio of acetaldehyde to methyl iodide in the feed liquid and the bottom residue. The organic phase (methyl iodide phase) obtained by separating the side stream may be recycled to the distillation column. In this case, the position of recycling the organic phase obtained by separating the side stream is preferably below the position of extracting the side stream, and preferably above the position of supplying the feed liquid, in the height direction of the column. Further, a miscible solvent for components (for example, methyl acetate and the like) constituting the organic phase separated from the process stream may be introduced into the distillation column (extractive distillation column). Examples of the miscible solvent include acetic acid and ethyl acetate. The introduction position of the miscible solvent is preferably below the side-stream drawing position, and more preferably above the feed position of the feed liquid, in the height direction of the column. In addition, when the organic phase obtained by separating the side stream is recycled to the distillation column, the introduction position of the miscible solvent is preferably below the recycling position. By recycling the organic phase obtained by separating the side stream to the distillation column or introducing the miscible solvent into the distillation column, the methyl acetate concentration in the extract extracted as the side stream can be reduced, and the methyl acetate concentration in the aqueous phase obtained by separating the extract can be reduced, whereby the contamination of methyl iodide into the aqueous phase can be suppressed.

The theoretical plates of the distillation column (extractive distillation column) are, for example, 1 to 100, preferably 2 to 50, more preferably 3 to 30, and particularly preferably 5 to 20 layers, and acetaldehyde can be efficiently separated and removed with a smaller number of layers as compared with 80 to 100 layers of conventional distillation columns and extractive distillation columns for removing acetaldehyde. The mass ratio of the flow rate of the extraction solvent to the flow rate of the feed liquid (organic phase and/or aqueous phase obtained by separating the process stream) (the former/latter) may be selected from the range of 0.0001/100 to 100/100, and is usually 0.0001/100 to 20/100, preferably 0.001/100 to 10/100, more preferably 0.01/100 to 8/100, and still more preferably 0.1/100 to 5/100. The column top temperature of the distillation column (extractive distillation column) is, for example, 15 to 120 ℃, preferably 20 to 90 ℃, more preferably 20 to 80 ℃, and still more preferably 25 to 70 ℃. The pressure at the top of the column is, for example, about 0.1 to 0.5MPa in absolute terms. Other conditions of the distillation column (extractive distillation column) may be the same as those of existing distillation columns and extractive distillation columns for removing acetaldehyde.

Fig. 3 is a schematic flow chart showing an example of the above-described acetaldehyde separation and removal system using extractive distillation. In this example, an organic phase and/or an aqueous phase (feed liquid) obtained by separating the process stream is supplied to a middle stage (a position between the top and bottom of the column) of the distillation column 94 through a supply line 201, water is introduced from the vicinity of the top of the column through a line 202, and extractive distillation is performed in the distillation column 94 (extractive distillation column). Above the feed point of the distillation column 94, a chimney tray 200 is disposed for receiving a liquid (extract) descending from a concentration region where methyl iodide and acetaldehyde in the column are concentrated. In the extractive distillation, it is preferable that the liquid on the chimney tray 200 is entirely extracted and introduced into the decanter 95 via the line 208 to be separated. The aqueous phase (containing acetaldehyde) in the decanter 95 is introduced into a cooler 95a through a line 212 and cooled, and the methyl iodide dissolved in the aqueous phase is separated into two phases, and the phases are separated by a decanter 96. The aqueous phase in the decanter 96 is fed to a distillation column 97 (acetaldehyde removal column) through a line 216 to be distilled, the vapor at the top of the column is introduced into a condenser 97a through a line 217 to be condensed, a part of the condensate (mainly acetaldehyde and methyl iodide) is refluxed to the top of the distillation column 97, and the remainder is either discarded or fed to a distillation column 98 (extractive distillation column) through a line 220. Water is introduced from the vicinity of the top of the distillation column 98 through a line 222, and extractive distillation is performed. The vapor at the top of the column is introduced into the condenser 98a through the line 223 to be condensed, and a part of the condensate (mainly methyl iodide) is refluxed to the top of the column, and the remainder is recycled to the reaction system through the line 226, but may be discharged outside the system. For the organic phase (methyl iodide phase) in decanter 95, it is preferred that the entire amount is recycled via

lines

209, 210 to below the location of chimney tray 200 of distillation column 94. A portion of the aqueous phase of decanter 95, and the organic phase of decanter 96, are recycled to distillation column 94 via

lines

213, 210,

lines

214, 210, respectively, but there are also instances where they are not recycled. A part of the aqueous phase of the decanter 95 can be used as an extraction solvent (water) in the distillation column 94. A portion of the aqueous phase of decanter 96 may be recycled to distillation column 94 via line 210. In some cases (for example, in the case where methyl acetate is contained in the feed liquid), a miscible solvent (for example, acetic acid, ethyl acetate, etc.) with respect to components (for example, methyl acetate, etc.) constituting an organic phase obtained by separating the process stream is fed to the distillation column 94 through the line 215, whereby the distillation efficiency can be improved. The position of the supply of the miscible solvent to the distillation column 94 is above the feed liquid supply portion (the connection portion of the line 201) and below the connection portion of the recirculation line 210. The bottom residue of the distillation column 94 is recycled to the reaction system. The vapor at the top of the distillation column 94 is introduced into a condenser 94a through a line 203 and condensed, the condensate is separated in a decanter 99, the organic phase is refluxed to the top of the distillation column 94 through a line 206, and the aqueous phase is introduced into a decanter 95 through a line 207. The bottom residue (water as a main component) of the distillation column 97 and the bottom residue (containing a small amount of acetaldehyde) of the distillation column 98 (extractive distillation column) are removed out of the system through

lines

218 and 224, respectively, or recycled to the reaction system. The gas that is not condensed in the

condensers

94a, 97a, 98a (

lines

211, 221, 227) is subjected to an absorption treatment or a disposal treatment in the scrubbing system 8.

Fig. 4 is a schematic flow chart showing another example of the above-described acetaldehyde separation and removal system using extractive distillation. In this example, the condensate of the vapor at the top of distillation column 94 is directed to holding tank 100, the entire amount of which is refluxed to the top of distillation column 94 via line 206. Otherwise, the same procedure as in the example of fig. 3 is followed.

FIG. 5 is a schematic flow diagram showing another example of the above-described acetaldehyde separation and removal system using extractive distillation. In this example, the liquid on the chimney tray 200 is extracted in full amount, and is directly introduced into the cooler 95a through the line 208 without passing through the decanter 95, cooled, and supplied to the decanter 96. Otherwise, the same procedure as in the example of fig. 4 is followed.

In fig. 1, the gas generated in the condenser 3a contains, for example, carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like, and is supplied from the condenser 3a to the scrubbing system 8 through lines 32 and 15. Methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like in the gas portion that reaches the scrubbing system 8 are absorbed by the absorbing liquid in the scrubbing system 8. The hydrogen iodide reacts with methanol or methyl acetate in the absorbing liquid to produce methyl iodide. Then, the liquid portion containing the useful components such as methyl iodide is recycled from the washing system 8 to the reaction tank 1 through the

recycle lines

48 and 23 and reused.

In the bottom liquid withdrawn from the bottom of the distillation column 3, a component having a boiling point higher than that of acetic acid (high boiling component) containing, for example, propionic acid and the above-mentioned catalyst and co-catalyst entrained by entrainment is contained more in the column bottom liquid than the above-mentioned overhead flow and side flow from the distillation column 3. The residue in the column still contains acetic acid, methyl iodide, methyl acetate, water, etc. In the present embodiment, a part of the bottom residual liquid is continuously introduced into the evaporation tank 2 through the

lines

25 and 26 and recycled, and the other part of the bottom residual liquid is continuously introduced into the reaction tank 1 through the

lines

25 and 23 and recycled.

In a preferred embodiment, the column bottom residual liquid extracted from the bottom of the distillation column 3 contains a component having a boiling point higher than that of acetic acid (high boiling point component) more than the above-mentioned overhead flow and side flow from the distillation column 3, and contains, for example, propionic acid, and the above-mentioned catalyst and co-catalyst entrained by entrainment. The bottom residue also contains acetic acid, methyl iodide, methyl acetate, crotonaldehyde, 2-ethyl crotonaldehyde, butyl acetate, water, and the like. In the present embodiment, a part of the bottom residual liquid is continuously introduced into the evaporation tank 2 through the

lines

25 and 23 and recycled.

The 1 st acetic acid stream, which is continuously withdrawn as a side stream from distillation column 3, is richer in acetic acid than the vapor stream that is continuously introduced into distillation column 3. That is, the acetic acid concentration of the 1 st acetic acid stream is higher than the acetic acid concentration of the vapor stream. The acetic acid concentration of the 1 st acetic acid stream is, for example, 90 to 99.9 mass%, preferably 93 to 99 mass%. In addition, the 1 st acetic acid stream may contain, for example, methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, acetaldehyde, formic acid, propionic acid, and the like, in addition to acetic acid. In the 1 st acetic acid stream, the methyl iodide concentration is, for example, 8 mass% or less (for example, 0.1 to 8 mass%), preferably 0.2 to 5 mass%, the water concentration is, for example, 8 mass% or less (for example, 0.1 to 8 mass%), preferably 0.2 to 5 mass%, and the methyl acetate concentration is, for example, 8 mass% or less (for example, 0.1 to 8 mass%), preferably 0.2 to 5 mass%. The connection position of the line 27 to the distillation column 3 may be, as shown in the figure, above the connection position of the line 21 to the distillation column 3 in the height direction of the distillation column 3, may be below the connection position of the line 21 to the distillation column 3, or may be the same as the connection position of the line 21 to the distillation column 3. The 1 st acetic acid stream from distillation column 3 is continuously introduced at a given flow rate through line 27 to the subsequent distillation column 5. The 1 st acetic acid stream extracted as a side stream of the distillation column 3, the bottom liquid of the distillation column 3, or the condensate of the vapor at the bottom of the distillation column 3 may be directly introduced as the product acetic acid, or may be directly and continuously introduced into the distillation column 6 without passing through the distillation column 5.

In a preferred embodiment, the 1 st acetic acid stream withdrawn continuously as a side stream from distillation column 3 is richer in acetic acid than the vapor stream that is continuously introduced into distillation column 3. That is, the acetic acid concentration of the 1 st acetic acid stream is higher than the acetic acid concentration of the vapor stream. The acetic acid concentration of the 1 st acetic acid stream is, for example, 90 to 99.9 mass%, preferably 93 to 99 mass%. The 1 st acetic acid stream may contain, in addition to acetic acid, methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, 2-ethylcrotonaldehyde, butyl acetate, formic acid and propionic acid, and iodoalkanes such as iodoethane, iodopropane, iodobutane, iodohexane and iododecane. In the 1 st acetic acid stream, the methyl iodide concentration is, for example, 0.1 to 8% by mass, preferably 0.2 to 5% by mass, the water concentration is, for example, 0.1 to 8% by mass, preferably 0.2 to 5% by mass, and the methyl acetate concentration is, for example, 0.1 to 8% by mass, preferably 0.2 to 5% by mass.

In the present invention, since the reflux ratio of the 1 st distillation column is set to a specific value or more, crotonaldehyde is concentrated to the top of the distillation column. Thus, the crotonaldehyde concentration in the 1 st acetic acid stream extracted as a side stream of the 1 st distillation column is low. The crotonaldehyde concentration in the 1 st acetic acid stream is, for example, 1.3 mass ppm or less, preferably 1.0 mass ppm or less, more preferably 0.85 mass ppm or less, and particularly preferably 0.5 mass ppm or less (for example, 0.25 mass ppm or less). By setting the crotonaldehyde concentration in the 1 st acetic acid stream to 1.3 ppm by mass or less, the crotonaldehyde concentration in the 2 nd acetic acid stream described later can be significantly reduced, and the potassium permanganate test value in the 2 nd acetic acid stream can be significantly increased. The lower limit of the crotonaldehyde concentration in the 1 st acetic acid stream may be 0 mass ppm, but may be, for example, 0.01 mass ppm (or 0.10 mass ppm). The concentration of 2-ethylcrotonaldehyde in the 1 st acetic acid stream is, for example, 1.0 mass ppm or less, preferably 0.50 mass ppm or less. By setting the concentration of 2-ethylcrotonaldehyde in the 1 st acetic acid stream to 1.0 mass ppm or less, the potassium permanganate test value of the 2 nd acetic acid stream described later can be further improved. The lower limit of the concentration of 2-ethylcrotonaldehyde in the 1 st acetic acid stream may be, for example, 0 mass ppm or 0.01 mass ppm (or 0.10 mass ppm). The butyl acetate concentration in the 1 st acetic acid stream is, for example, 15 mass ppm or less, preferably 10 mass ppm or less, more preferably 8 mass ppm or less, and particularly preferably 5 mass ppm or less (for example, 3 mass ppm or less). By setting the butyl acetate concentration in the 1 st acetic acid stream to 15 mass ppm or less, the purity of the 2 nd acetic acid stream described later can be improved. The lower limit of the butyl acetate concentration in the 1 st acetic acid stream may be, for example, 0 mass ppm, or 0.1 mass ppm (e.g., 0.3 mass ppm or 1.0 mass ppm).

In the preferred embodiment, the connection position of the line 27 to the distillation column 3 may be, as shown in the drawing, above the connection position of the line 21 to the distillation column 3 in the height direction of the distillation column 3, may be below the connection position of the line 21 to the distillation column 3, or may be the same as the connection position of the line 21 to the distillation column 3. The 1 st acetic acid stream from distillation column 3 is continuously introduced at a given flow rate through line 27 to the subsequent distillation column 5.

In the preferred embodiment, as long as the quality can be tolerated, the bottom residue extracted from the bottom of the distillation column 3 or the 1 st acetic acid stream extracted as a side stream from the distillation column 3 may be used as the product acetic acid as it is.

In the 1 st acetic acid stream flowing through line 27, potassium hydroxide may be supplied or added via line 55 (potassium hydroxide introduction line). The potassium hydroxide may be supplied or added in the form of a solution such as an aqueous solution. By supplying or adding potassium hydroxide to the 1 st acetic acid stream, hydrogen iodide in the 1 st acetic acid stream can be reduced. Specifically, hydrogen iodide reacts with potassium hydroxide to produce potassium iodide and water. This can reduce corrosion of the apparatus such as the distillation column due to hydrogen iodide. Potassium hydroxide may be supplied or added to a suitable place in the process where hydrogen iodide is present. Potassium hydroxide added in the process was also reacted with acetic acid to produce potassium acetate.

The distillation column 5 is a unit for performing the 2 nd distillation step, and may be regarded as a so-called dehydration column in the present embodiment. The 2 nd distillation step is a step for further purifying acetic acid by subjecting the 1 st acetic acid stream continuously introduced into the distillation column 5 to distillation treatment. The material of the distillation column 5 (at least the material of the liquid contact portion and the gas contact portion) is preferably a nickel-based alloy or zirconium. By using such a material, corrosion of the inside of the distillation column by hydrogen iodide and acetic acid can be suppressed, and elution of corrosion metal ions can be suppressed.

The feed to distillation column 5 comprises at least a portion of the 1 st acetic acid stream (line 27), and streams other than the 1 st acetic acid stream [ e.g., a recycle stream from a downstream process (e.g., line 42) ] can also be added.

The distillation column 5 includes a rectifying column such as a plate column and a packed column. When a tray column is used as the distillation column 5, the theoretical plate number is, for example, 5 to 50 stages. The reflux ratio of the distillation column 5 is, for example, 0.1 or more, preferably 0.3 or more (for example, 0.32 or more), more preferably 1.0 or more, further preferably 5.0 or more, and particularly preferably 10 or more (for example, 12 or more). The upper limit of the reflux ratio of the distillation column 5 may be, for example, about 3000 (or 1000) or about 200 (or 100). When the reflux ratio of the distillation column 5 is 0.1 or more, since crotonaldehyde has a lower boiling point than acetic acid, crotonaldehyde flowing into the distillation column 5 can be concentrated at the top of the column, and the concentration of crotonaldehyde in the 2 nd acetic acid stream obtained as a side stream or a column bottom stream can be remarkably reduced. In addition, if the overhead stream (2 nd overhead stream) of the distillation column 5 overhead in which crotonaldehyde is concentrated is recycled to the reaction tank 1, as described above, crotonaldehyde can be converted into 2-ethylcrotonaldehyde which is less harmful to the potassium permanganate test value and harmless butyl acetate, and therefore the quality of acetic acid is further improved.

The column top pressure in the distillation column 5 existing in the 2 nd distillation step is, for example, 0.01 to 0.50MPa (gauge pressure), preferably 0.10 to 0.28MPa (gauge pressure), more preferably 0.15 to 0.23MPa (gauge pressure), and still more preferably 0.17 to 0.21MPa (gauge pressure). The column bottom pressure is higher than the column top pressure, and is, for example, 0.13 to 0.31MPa (gauge pressure), preferably 0.18 to 0.26MPa (gauge pressure), and more preferably 0.20 to 0.24MPa (gauge pressure). In the distillation column 5 existing in the 2 nd distillation step, it is preferable that the temperature at the top of the column is lower than 175 ℃ (preferably lower than 165 ℃) and the temperature at the bottom of the column is lower than 185 ℃ (preferably lower than 175 ℃). When the column top temperature and the column bottom temperature of the distillation column 5 are set to the above ranges, corrosion of the inside of the distillation column by hydrogen iodide and acetic acid can be further suppressed, and elution of corrosion metal ions can be further suppressed. The overhead temperature is more preferably less than 163 ℃, still more preferably less than 161 ℃, particularly preferably less than 160 ℃, and especially preferably less than 155 ℃. The lower limit of the overhead temperature is, for example, 110 ℃. The column bottom temperature is more preferably lower than 173 ℃, still more preferably lower than 171 ℃, particularly preferably lower than 166 ℃. The lower limit of the bottom temperature is, for example, 120 ℃.

In the preferred embodiment, the distillation column 5 is a unit for performing the 2 nd distillation step, and may be referred to as a so-called dehydration column in the present embodiment. The 2 nd distillation step is a step for further purifying acetic acid by distilling the 1 st acetic acid stream continuously introduced into the distillation column 5. The distillation column 5 includes a rectifying column such as a plate column and a packed column. When a plate column is used as the distillation column 5, the number of theoretical plates is, for example, 5 to 50, and the reflux ratio is, for example, 0.2 to 3000 depending on the number of theoretical plates. The column top pressure in the distillation column 5 existing in the 2 nd distillation step is set to 150 to 250kPa (gauge pressure), for example, and the column bottom pressure is set to be higher than the column top pressure, for example, 160 to 290kPa (gauge pressure). The column top temperature in the distillation column 5 existing in the 2 nd distillation step is set to, for example, a temperature of 130 to 160 ℃ higher than the boiling point of water and lower than the boiling point of acetic acid at a set column top pressure, and the column bottom temperature is set to, for example, a temperature of 150 to 175 ℃ higher than the boiling point of acetic acid at a set column bottom pressure.

From the column top of the distillation column 5, a vapor as an overhead stream (2 nd overhead stream) is continuously extracted to the line 33. From the bottom of the distillation column 5, a still bottoms liquid is continuously withdrawn to line 34. And 5b is a reboiler. From a height position between the top and bottom of the column in the distillation column 5, a side stream (liquid or gas) may also be continuously withdrawn to the line 34.

In the vapor withdrawn from the top of the distillation column 5, components having a boiling point lower than that of acetic acid (low boiling components) are contained more than the above-mentioned bottom residue from the distillation column 5, including, for example: methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like. Such vapor is continuously introduced into the condenser 5a through the line 33.

In a preferred embodiment, in the vapor withdrawn from the column top of the distillation column 5, components having a boiling point lower than that of acetic acid (low boiling components) are contained more than the above-mentioned still bottoms from the distillation column 5, including, for example: methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, crotonaldehyde, formic acid and the like. Such vapor is continuously introduced into the condenser 5a through a line 33.

The condenser 5a partially condenses the vapor from the distillation column 5 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components include, for example, water and acetic acid. A part of the condensed components is continuously refluxed from the condenser 5a to the distillation column 5 through a line 35. The other part of the condensed component is continuously introduced from the condenser 5a into the reaction tank 1 through the

lines

35, 36, and 23 and is recycled. The gas component generated in the condenser 5a includes, for example: carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like are supplied from the condenser 5a to the washing system 8 through the lines 37, 15. The hydrogen iodide in the gas component reaching the scrubbing system 8 is absorbed by the absorbing liquid in the scrubbing system 8, and methyl iodide is produced by the reaction between the hydrogen iodide in the absorbing liquid and methanol or methyl acetate, and then the liquid component containing the useful component such as methyl iodide is recycled from the scrubbing system 8 to the reaction tank 1 through the

recycling lines

48 and 23 and reused.

The column bottom residual liquid (or side stream) extracted from the bottom of the distillation column 5 contains more components (high boiling point components) having a boiling point higher than that of acetic acid than the column top stream from the distillation column 5, including, for example, propionic acid, potassium acetate (in the case where potassium hydroxide is supplied to the line 27 or the like), and the catalyst, promoter, and the like entrained in the mist. The bottoms from the column may also contain acetic acid. This bottoms liquid is continuously introduced into the subsequent distillation column 6 via line 34 as the 2 nd acetic acid stream.

The 2 nd acetic acid stream is richer in acetic acid than the 1 st acetic acid stream that is continuously introduced into distillation column 5. That is, the acetic acid concentration of the 2 nd acetic acid stream is higher than the acetic acid concentration of the 1 st acetic acid stream. The acetic acid concentration of the 2 nd acetic acid stream is in a range higher than that of the 1 st acetic acid stream, for example, 99.1 to 99.99 mass%. As described above, the 2 nd acetic acid stream contains, for example, propionic acid, hydrogen iodide, and the like in addition to acetic acid. In the present embodiment, in the case of extracting a side stream, the extraction position of the side stream from the distillation column 5 is lower than the introduction position of the 1 st acetic acid stream into the distillation column 5 in the height direction of the distillation column 5.

In a preferred embodiment, the bottom liquid withdrawn from the bottom of the distillation column 5 or a side stream (2 nd acetic acid stream) withdrawn from an intermediate position of the column is richer in acetic acid than the 1 st acetic acid stream continuously introduced into the distillation column 5. That is, the acetic acid concentration of the 2 nd acetic acid stream is higher than the acetic acid concentration of the 1 st acetic acid stream. The acetic acid concentration of the 2 nd acetic acid stream is in a range higher than that of the 1 st acetic acid stream, for example, 99.1 to 99.99 mass%. In the present embodiment, in the case of extracting a side stream, the extraction position of the side stream from the distillation column 5 is lower than the introduction position of the 1 st acetic acid stream into the distillation column 5 in the height direction of the distillation column 5.

In the present invention, the 2 nd acetic acid stream can be used directly as product acetic acid due to its high potassium permanganate test value. But may contain a trace amount of impurities [ for example, crotonaldehyde, 2-ethylcrotonaldehyde, butyl acetate, propionic acid, potassium acetate (in the case where potassium hydroxide is supplied to the line 27 or the like), hydrogen iodide, and the above-mentioned catalyst and co-catalyst entrained in the mist ]. Thus, the bottoms or side stream can be continuously introduced via line 34 into distillation column 6 for distillation.

The crotonaldehyde concentration in the 2 nd acetic acid stream is, for example, 0.98 mass ppm or less, preferably 0.80 mass ppm or less, more preferably 0.50 mass ppm or less, further preferably 0.30 mass ppm or less, and particularly preferably 0.17 mass ppm or less. By setting the crotonaldehyde concentration in the 2 nd acetic acid stream to 0.98 mass ppm or less, the test value for potassium permanganate in the 2 nd acetic acid stream can be greatly increased while the crotonaldehyde concentration in the 2 nd acetic acid stream is significantly reduced. The lower limit of the crotonaldehyde concentration in the 2 nd acetic acid stream may also be 0 mass ppm, but may also be, for example, 0.01 mass ppm (or 0.10 mass ppm). The 2-ethylcrotonaldehyde concentration in the 2 nd acetic acid stream is, for example, 1.0 mass ppm or less, preferably 0.50 mass ppm or less, more preferably 0.30 mass ppm or less, and further preferably 0.20 mass ppm or less. The potassium permanganate test value of the 2 nd acetic acid stream can be further improved by adjusting the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream to 1.0 mass ppm or less. The lower limit of the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream may be, for example, 0 mass ppm or 0.01 mass ppm (e.g., 0.10 mass ppm).

The butyl acetate concentration in the 2 nd acetic acid stream is, for example, 15 mass ppm or less, preferably 10 mass ppm or less, more preferably 8 mass ppm or less, and particularly preferably 5 mass ppm or less (for example, 3 mass ppm or less). The purity of the 2 nd acetic acid stream can be increased by setting the butyl acetate concentration in the 2 nd acetic acid stream to 15 mass ppm or less. The lower limit of the butyl acetate concentration in the 2 nd acetic acid stream may be, for example, 0 mass ppm, or 0.1 mass ppm (e.g., 0.3 mass ppm or 1.0 mass ppm).

The potassium permanganate test value of the 2 nd acetic acid stream is preferably more than 50 minutes, more preferably 60 minutes or more, further preferably 100 minutes or more, and particularly preferably 120 minutes or more (for example, 180 minutes or more, particularly 240 minutes or more, and more preferably 360 minutes or more).

In the 2 nd acetic acid stream flowing through line 34, potassium hydroxide may be supplied or added via line 56 (potassium hydroxide introduction line). The potassium hydroxide may be supplied or added in the form of a solution such as an aqueous solution. By supplying or adding potassium hydroxide to the 2 nd acetic acid stream, hydrogen iodide in the 2 nd acetic acid stream can be reduced. Specifically, hydrogen iodide reacts with potassium hydroxide to produce potassium iodide and water. This can reduce corrosion of the apparatus such as the distillation column due to hydrogen iodide.

The distillation column 6 is a unit for performing the 3 rd distillation step, and is regarded as a so-called high boiling point removal column in the present embodiment. The 3 rd distillation step is a step for purifying acetic acid by purifying the 2 nd acetic acid stream continuously introduced into the distillation column 6. In this embodiment, the step is not necessarily essential. The distillation column 6 includes, for example, a rectifying column such as a plate column and a packed column. When a plate column is used as the distillation column 6, the number of theoretical plates is, for example, 5 to 50, and the reflux ratio is, for example, 0.2 to 3000 depending on the number of theoretical plates. The column top pressure in the distillation column 6 existing in the 3 rd distillation step is set to, for example, -100 to 150kPa (gauge pressure), and the column bottom pressure is set to be higher than the column top pressure, for example, -90 to 180kPa (gauge pressure). The column top temperature in the distillation column 6 existing in the 3 rd distillation step is set to a temperature of, for example, 50 to 150 ℃ higher than the boiling point of water at a set column top pressure and lower than the boiling point of acetic acid, and the column bottom temperature is set to a temperature of, for example, 70 to 160 ℃ higher than the boiling point of acetic acid at a set column bottom pressure.

From the top of the distillation column 6, vapor is continuously withdrawn as an overhead stream to line 38. From the bottom of the distillation column 6, a still bottoms liquid is continuously withdrawn to line 39.6b is a reboiler. From a height position between the top and bottom of the column in the distillation column 6, a side stream (liquid or gas) is continuously withdrawn to the line 46. The connection position of the line 46 to the distillation column 6 in the height direction of the distillation column 6 may be above the connection position of the line 34 to the distillation column 6 as shown in the figure, may be below the connection position of the line 34 to the distillation column 6, or may be the same as the connection position of the line 34 to the distillation column 6.

The vapor withdrawn from the top of the distillation column 6 contains a component having a boiling point lower than that of acetic acid (low boiling point component) more than the above-mentioned bottom residue from the distillation column 6, and includes, for example, methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, formic acid, and the like in addition to acetic acid. Such vapor is continuously introduced into the condenser 6a through a line 38.

The condenser 6a partially condenses the vapor from the distillation column 6 by cooling it, thereby separating it into a condensed component and a gas component. The condensed components include, for example, methyl iodide, hydrogen iodide, water, methyl acetate, dimethyl ether, methanol, formic acid, and the like, in addition to acetic acid. At least a part of the condensed components is continuously refluxed from the condenser 6a to the distillation column 6 through the line 40. Part of the condensed components (distillate components) can be recycled from the condenser 6a via

lines

40, 41, 42 to the 1 st acetic acid stream in the line 27 before introduction to the distillation column 5. Simultaneously with this or instead, a part of the condensed portion (distillate components) may be recycled from condenser 6a via

lines

40, 41, 43 to the vapor stream in line 21 before being introduced into distillation column 3. Further, as for a part of the condensed portion (distillate component), it may be recycled from the condenser 6a to the reaction tank 1 through the

lines

40, 44, 23. Further, a part of the distillate component from the condenser 6a may be supplied to the washing system 8 as described above and used as an absorbent in the system. In the washing system 8, the gas component having absorbed the useful component is discharged to the outside of the apparatus, and further, the liquid component containing the useful component is introduced or recycled from the washing system 8 to the reaction tank 1 through the

recycle lines

48 and 23 to be reused. A part of the distillate component from the condenser 6a may be guided to various pumps (not shown) operating in the apparatus through lines other than those shown in the drawings and used as a sealing liquid for the pumps. Further, a part of the distillate component from the condenser 6a may be extracted outside the apparatus through an extraction line provided in the line 40 at regular intervals, or may be extracted outside the apparatus in an unusual manner as necessary. When a part of the condensed portion (distillate component) is removed from the distillation treatment system of the distillation column 6, the amount of the distillate component (distillate amount) is, for example, 0.01 to 30% by mass, preferably 0.1 to 10% by mass, more preferably 0.3 to 5% by mass, and still more preferably 0.5 to 3% by mass of the condensate produced in the condenser 6a. On the other hand, the gas components generated in the condenser 6a include, for example, carbon monoxide, hydrogen, methane, carbon dioxide, nitrogen, oxygen, methyl iodide, hydrogen iodide, water, methyl acetate, acetic acid, dimethyl ether, methanol, acetaldehyde, formic acid, and the like, and are supplied from the condenser 6a to the scrubbing system 8 through the

lines

45, 15.

The bottom residual liquid extracted from the bottom of the distillation column 6 through the line 39 contains more components having a boiling point higher than that of acetic acid (high boiling point components) including, for example, propionic acid, potassium acetate (in the case where potassium hydroxide is supplied to the line 34 and the like) and the like than the overhead flow from the distillation column 6. The bottom residue extracted from the bottom of the distillation column 6 through the line 39 also contains corrosive metals such as metals that are generated and released on the inner walls of the components of the acetic acid production apparatus, and compounds of iodine derived from corrosive iodine and the corrosive metals. Such a bottom residue is discharged to the outside of the acetic acid production apparatus in the present embodiment.

In a preferred embodiment, the bottom liquid extracted from the bottom of the distillation column 6 through the line 39 contains more components having a boiling point higher than that of acetic acid (high boiling point components) including, for example, acetic acid salts such as propionic acid and potassium acetate (in the case where an alkali such as potassium hydroxide is supplied to the line 34) and the like than the overhead flow from the distillation column 6. The bottom residual liquid extracted from the bottom of the distillation column 6 through the line 39 also contains corrosive metals such as metals generated and released on the inner walls of the components of the acetic acid production apparatus, and compounds of iodine derived from corrosive iodine and the corrosive metals. Such a bottom residue is discharged to the outside of the acetic acid production apparatus in the present embodiment.

The side stream continuously withdrawn from the distillation column 6 to the line 46 is continuously introduced into the subsequent ion exchange resin column 7 as the 3 rd acetic acid stream. This 3 rd acetic acid stream is richer in acetic acid than the 2 nd acetic acid stream that is continuously introduced into the distillation column 6. That is, the acetic acid concentration of the 3 rd acetic acid stream is higher than the acetic acid concentration of the 2 nd acetic acid stream. The acetic acid concentration of the 3 rd acetic acid stream is in a range higher than that of the 2 nd acetic acid stream, for example, 99.8 to 99.999 mass%. In the present embodiment, the position of the side stream withdrawn from the distillation column 6 is higher than the position of the 2 nd acetic acid introduced into the distillation column 6 in the height direction of the distillation column 6. In other embodiments, the position of the side stream taken from the distillation column 6 is lower than or the same as the position of the introduction of the 2 nd acetic acid stream into the distillation column 6 in the height direction of the distillation column 6. The distillation column 6 may be replaced with a single-type distiller (evaporator), and the distillation column 6 may be omitted if impurities can be sufficiently removed by the distillation column 5.

In a preferred embodiment, the side stream continuously withdrawn from distillation column 6 to line 46 is continuously introduced as the 3 rd acetic acid stream to the subsequent ion exchange resin column 7. This 3 rd acetic acid stream is richer in acetic acid than the 2 nd acetic acid stream that is continuously introduced into distillation column 6. That is, the acetic acid concentration of the 3 rd acetic acid stream is higher than the acetic acid concentration of the 2 nd acetic acid stream. The acetic acid concentration of the 3 rd acetic acid stream is in a range higher than that of the 2 nd acetic acid stream, for example, 99.8 to 99.999 mass%. In the present embodiment, the position of the side stream withdrawn from the distillation column 6 is higher than the position of the 2 nd acetic acid stream introduced into the distillation column 6 in the height direction of the distillation column 6. In other embodiments, the position of the side stream taken from the distillation column 6 is lower than or the same as the position of the introduction of the 2 nd acetic acid stream into the distillation column 6 in the height direction of the distillation column 6. It should be noted that the distillation column 6 may be replaced with a single-type distiller (evaporator). In particular, in the present invention, since acetic acid having a very high potassium permanganate test value can be obtained by the distillation treatment in the distillation column 5, the distillation column 6 can be omitted.

The ion exchange resin column 7 is a purification unit for performing an adsorption removal step. This adsorption removal step is a step for adsorbing and removing mainly alkyl iodide (such as ethyl iodide and decyl iodide) contained in a trace amount in the 3 rd acetic acid stream continuously introduced into the ion exchange resin column 7, thereby further purifying the acetic acid. In the ion exchange resin column 7, an ion exchange resin having an adsorption ability to iodoalkane is filled in the column to form an ion exchange resin bed. Examples of such an ion exchange resin include: cation exchange resins in which a part of leaving proton in a sulfonic acid group, a carboxyl group, a phosphonic acid group, or the like as an exchange group is replaced with a metal such as silver or copper. In the adsorption removal step, for example, the 3 rd acetic acid stream (liquid) is passed through an ion exchange resin column 7 packed with such an ion exchange resin, and impurities such as alkyl iodide in the 3 rd acetic acid stream are adsorbed to the ion exchange resin and removed from the 3 rd acetic acid stream in the passing process. In the ion exchange resin column 7 existing in the adsorption removal step, the internal temperature is, for example, 18 to 100 ℃, and the flow rate of acetic acid [ resin volume per 1m ] ³ Acetic acid treatment amount (m) ³ /h)]Is, for example, 3 to 15m ³ /h·m ³ (resin volume).

In a preferred embodiment, the ion exchange resin column 7 is a purification unit for performing an adsorption removal step. This adsorption removal step is a step for adsorbing and removing mainly alkyl iodide (for example, ethyl iodide, propyl iodide, butyl iodide, hexane iodide, decane iodide, etc.) contained in a trace amount in the 3 rd acetic acid stream continuously introduced into the ion exchange resin column 7 to further purify the acetic acid. The distillation column 6 may be omitted, and the 2 nd acetic acid stream from the distillation column 5 may be supplied to the ion exchange resin column 7. Further, the adsorption and removal step using the ion exchange resin column 7 may not be provided.

In the preferred embodiment, in the ion exchange resin column 7, an ion exchange resin having an adsorption ability to iodoalkane is filled in the column to form an ion exchange resin bed. Examples of such an ion exchange resin include: cation exchange resins in which a part of leaving proton in a sulfonic acid group, a carboxyl group, a phosphonic acid group, or the like as an exchange group is replaced with a metal such as silver or copper. In the adsorption removal step, for example, the 3 rd acetic acid stream (liquid) is passed through the ion exchange resin column 7 packed with such an ion exchange resin, and impurities such as alkyl iodide in the 3 rd acetic acid stream are adsorbed to the ion exchange resin and removed from the 3 rd acetic acid stream in the passing process. In the ion exchange resin column 7 existing in the adsorption removal step, the internal temperature is, for example, 18 to 100 ℃ and the flow rate of acetic acid [ resin volume per 1m ] ³ Acetic acid treatment amount (m) ³ /h)]Is, for example, 3 to 15m ³ /h·m ³ (resin volume).

A4 th acetic acid stream is continuously withdrawn from the lower end of the ion exchange resin column 7 into line 47. The acetic acid concentration of the 4 th acetic acid stream is higher than the acetic acid concentration of the 3 rd acetic acid stream. That is, the 4 th acetic acid stream is richer in acetic acid than the 3 rd acetic acid stream that is continuously introduced into the ion exchange resin column 7. The acetic acid concentration of the 4 th acetic acid stream is in a range higher than that of the 3 rd acetic acid stream, for example, 99.9 to 99.999 mass% or more. In the present production process, the 4 th acetic acid stream may be stored in a product tank outside the figure.

In this acetic acid production apparatus, as a purification unit for further purifying the above-mentioned 4 th acetic acid stream from the ion exchange resin column 7, a so-called product column or a finishing column as a distillation column may also be provided. When such a product column is provided, the product column includes, for example, a rectifying column such as a plate column and a packed column. When a tray column is used as the product column, the number of theoretical plates is, for example, 5 to 50 layers, and the reflux ratio is, for example, 0.5 to 3000 depending on the number of theoretical plates. The column top pressure in the product column in the purification step is set to, for example, -195 to 150kPa (gauge pressure), and the column bottom pressure is set to be higher than the column top pressure, for example, -190 to 180kPa (gauge pressure). The column top temperature is set to a temperature of, for example, 50 to 150 ℃ higher than the boiling point of water and lower than the boiling point of acetic acid at a set column top pressure, and the column bottom temperature is set to a temperature of, for example, 70 to 160 ℃ higher than the boiling point of acetic acid at a set column bottom pressure, inside the product column. It should be noted that the product column or the finishing column may be replaced with a single-type distiller (evaporator).

In the case where a product column is provided, all or a part of the 4 th acetic acid stream (liquid) from the ion exchange resin column 7 is continuously introduced into the product column. From the column top of such a product column, a vapor as an overhead stream containing a slight amount of low boiling components (for example, methyl iodide, water, methyl acetate, dimethyl ether, crotonaldehyde, acetaldehyde, formic acid, and the like) is continuously extracted. The vapor is separated into a condensed component and a gaseous component by a given condenser. Part of the condensed components is continuously refluxed to the product column, and the other part of the condensed components is recycled to the reaction tank 1, or discarded to the outside of the system, or both, and the gas part is supplied to the scrubbing system 8. From the bottom of the product column, a bottom residue containing a trace amount of high boiling components is continuously extracted, which is recycled to, for example, the 2 nd acetic acid stream in line 34 before being introduced into the distillation column 6. From a height position between the top and bottom of the column in the product column, a side stream (liquid) is continuously withdrawn as a 5 th acetic acid stream. The position of the side stream taken from the product column is lower in the height direction of the product column than the position of the introduction of the 4 th acetic acid stream into the product column, for example. With respect to the 5 th acetic acid stream, the acetic acid is more enriched than the 4 th acetic acid stream that is continuously introduced to the product column. That is, the acetic acid concentration of the 5 th acetic acid stream is higher than the acetic acid concentration of the 4 th acetic acid stream. The acetic acid concentration of the 5 th acetic acid stream is in a range higher than that of the 4 th acetic acid stream, for example, 99.9 to 99.999 mass% or more. The 5 th acetic acid stream may be stored, for example, in a product tank outside the figure. As an alternative to being disposed downstream of the distillation column 6 (or in addition to being disposed downstream of the distillation column 6), the ion exchange resin column 7 may be disposed downstream of the product column to treat the acetic acid stream from the product column.

Examples

The present invention will be described in more detail with reference to examples, but the present invention is not limited to these examples. The parts,%, ppm and ppb are all by mass. The water concentration was measured by the Karl Fischer (Karl Fischer) water content method, the metal ion concentration was measured by ICP analysis (or atomic absorption spectrometry), the potassium permanganate test value was measured by the procedure of the visual colorimetry according to JIS K1351 (1993 version), and the concentrations of other components were measured by gas chromatography. In addition, "AD" represents acetaldehyde, "MeI" represents methyl iodide, "MA" represents methyl acetate, and "AC" represents acetic acid. The AD distribution ratios in the aqueous phase and the organic phase and the MA distribution ratios in the aqueous phase and the organic phase were determined by the following equations.

First, the following experiment was performed to investigate conditions of the liquid separation step.

Experimental examples 1 to 10

Experimental examples 1 to 10 are experiments in which the concentration of each component in an aqueous phase and an organic phase and the distribution ratio of AD were examined by changing the AD concentration in a feed liquid supplied to a liquid separation step (liquid supplied to the liquid separation step) using a liquid-liquid equilibrium measuring apparatus 300 corresponding to the pressurized system shown in fig. 6. In fig. 6, 301 is a pressure-resistant glass vessel (internal volume 100 cc), 302 is a stirrer (rugby ball), 303 is a lower phase extraction tube, 304 is an upper phase extraction tube, 305 is a thermometer, 306 is a magnetic stirrer, 307 is a water bath, 308 is a temperature regulator, 309 is a thermometer, 310 is a lower phase sampling tube, 311 is an upper phase sampling tube, 312 is a pressure gauge, 313 is a pressure regulating valve, 314, 315, 316, and 317 are valves, 318 is a safety valve, 319 is a lower phase extraction line, 320 is an upper phase extraction line, 321 is a nitrogen introduction line, 322 is a pressure release line, and 323 is an exhaust line. The dashed line indicates the liquid level or interface.

AD, meI and water were charged into a pressure-resistant glass vessel 301 in amounts shown in the column of "actual feed amount" in table 1, a stirrer 302 was placed and the lid was closed, the inside was replaced with nitrogen, and after that, the temperature shown in the column of "temperature" in table 1 was adjusted by a water bath 307, and stirring was carried out at 300rpm for 30 minutes. After the stirring was stopped, the mixture was left for 10 minutes to completely separate the aqueous phase and the organic phase (methyl iodide phase) and then the aqueous phase and the organic phase were sampled by sampling tubes 311 and 310, respectively, and the concentrations of AD, meI, and water were measured to determine the AD distribution ratio. The results are shown in Table 1.

Experimental examples 11 to 15

In experimental examples 11 to 15, the AD concentration in the feed liquid was kept constant, and the concentrations of the respective components in the aqueous phase and the organic phase and the AD distribution ratio were examined while changing the temperature (liquid temperature) in the liquid separation step.

The AD distribution ratio was determined by measuring the concentrations of the respective components in the aqueous phase and the organic phase in the same manner as in experimental examples 1 to 10, with the feed amounts and temperatures of AD, meI and water to the pressure-resistant glass vessel 301 set to the values shown in table 1. The results are shown in Table 1.

Examples 16 to 21

In experimental examples 16 to 21, the AD concentration and MA concentration in the feed liquid were kept constant, and the concentrations of the respective components in the aqueous phase and the organic phase, the AD distribution ratio, and the MA distribution ratio were examined while changing the temperature (liquid temperature) in the liquid separation step.

The AD distribution ratio and MA distribution ratio were determined by measuring the concentrations of the respective components in the aqueous phase and the organic phase in the same manner as in experimental examples 1 to 10, with the feed amounts and temperatures of AD, meI, MA and water to be charged into the pressure-resistant glass vessel 301 set to the values shown in table 2. The results are shown in Table 2.

Experimental examples 22 to 26

Examples 22 to 26 are experiments in which the concentrations of the respective components in the aqueous phase and the organic phase, the AD distribution ratio, and the MA distribution ratio were examined by changing the MA concentration in the feed liquid.

Reference examples 1 to 2

Reference examples 1 to 2 are experiments in which the concentrations of the respective components in the aqueous phase and the organic phase and the AD distribution ratio were examined by changing the AC concentration in the feed liquid.

The AD distribution ratio was determined by measuring the concentrations of the respective components in the aqueous phase and the organic phase in the same manner as in experimental examples 1 to 10, with the feed amounts and temperatures of AD, meI, water and AC to be charged into the pressure-resistant glass vessel 301 set to the values shown in table 2. The results are shown in Table 2.

[ Table 1]

[ Table 2]

[ examination of results ]

From the results of experimental examples 1 to 10, the following tendency was found: the higher the AD concentration in the feed liquid, the higher the AD concentration in the aqueous phase and the organic phase, the smaller the AD distribution ratio, and the higher the distribution ratio of AD in the organic phase. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the AD concentration of the aqueous phase or organic phase is high, the advantage of the AD removal treatment of the organic phase is large, and when the AD concentration of the aqueous phase or organic phase is low, the advantage of the AD removal treatment of the organic phase is small. On the other hand, when the AD concentrations of the aqueous phase and the organic phase are low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, and when the AD concentrations of the aqueous phase and the organic phase are high, the advantage of subjecting the aqueous phase to the AD removal treatment is small. From another viewpoint, when a step or an apparatus for subjecting the organic phase to the AD removal treatment is used, it is preferable to increase the AD concentration of the aqueous phase and/or the organic phase in the liquid separation step by adjusting, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, and the like. In addition, when a step or an apparatus for subjecting the aqueous phase to the AD removal treatment is used, it is preferable to reduce the AD concentration of the aqueous phase and/or the organic phase in the liquid separation step by adjusting, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, and the like.

From the results of experimental examples 11 to 15, the following tendency was observed: under the condition that the AD concentration is constant, the higher the temperature at the time of liquid separation, the smaller the AD distribution ratio, and the higher the distribution ratio of AD in the organic phase. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the temperature at the time of liquid separation is high, the advantage of subjecting the organic phase to the AD removal treatment is large, and when the temperature at the time of liquid separation is low, the advantage of subjecting the organic phase to the AD removal treatment is small. On the other hand, when the temperature at the time of liquid separation is low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, and when the temperature at the time of liquid separation is high, the advantage of subjecting the aqueous phase to the AD removal treatment is small.

From the results of experimental examples 16 to 21, the following tendency was found: under the condition that the AD concentration and MA concentration are constant, the distribution ratio of AD is smaller as the temperature at the time of liquid separation is higher, and the distribution ratio of AD in the organic phase is relatively higher as the temperature is higher even in a system in which methyl acetate is present. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the temperature at the time of liquid separation is high, the advantage of performing the AD removal treatment on the organic phase is large, and when the temperature at the time of liquid separation is low, the advantage of performing the AD removal treatment on the organic phase is small. On the other hand, even in the system in which methyl acetate is present, when the temperature at the time of liquid separation is low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, and when the temperature at the time of liquid separation is high, the advantage of subjecting the aqueous phase to the AD removal treatment is small. From another point of view, when a step or a facility for subjecting the organic phase to the AD removal treatment is used regardless of the presence or absence of methyl acetate, it is preferable to set the liquid temperature at the time of liquid separation in the liquid separation step to a high value. In addition, when a step or a facility for subjecting the aqueous phase to the AD removal treatment is employed, the liquid temperature at the time of liquid separation in the liquid separation step is preferably set to a low value.

From the results of experimental examples 22 to 26 and experimental examples 4 and 17, the following tendency was observed: the higher the MA concentration in the feed liquid, the higher the MA concentration in the aqueous phase and the organic phase, the smaller the AD distribution ratio, and the higher the distribution ratio of AD in the organic phase. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the MA concentration of the aqueous phase or organic phase is high, the advantage of subjecting the organic phase to the AD removal treatment is large, and when the MA concentration of the aqueous phase or organic phase is low, the advantage of subjecting the organic phase to the AD removal treatment is small. On the other hand, when the MA concentration of the aqueous phase or the organic phase is low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, and when the MA concentration of the aqueous phase or the organic phase is high, the advantage of subjecting the aqueous phase to the AD removal treatment is small. On the other hand, when a step or an apparatus for subjecting the organic phase to the AD removal treatment is used, it is preferable to increase the MA concentration of the aqueous phase and/or the organic phase in the liquid separation step by adjusting, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, and the like. In addition, when a step or an apparatus for subjecting the aqueous phase to the AD removal treatment is used, it is preferable to reduce the MA concentration of the aqueous phase and/or the organic phase in the liquid separation step by adjusting, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, and the like. From the results of experimental examples 22 to 26, it was found that the higher the MA concentration in the feed liquid, the higher the MA concentration in the aqueous phase and the organic phase, and the smaller the MA distribution ratio as well as the AD distribution ratio.

From the results of reference examples 1 and 2, it was found that the AD distribution ratio did not change much depending on the AC concentrations of the feed liquid, the aqueous phase, and the organic phase. That is, the AC concentration is not significant in comparing the advantages of the method of subjecting the organic phase to the AD removal treatment and the method of subjecting the aqueous phase to the AD removal treatment from the viewpoint of the removal efficiency of acetaldehyde.

Next, the following description will be given by taking examples as examples. Although the present example is based on a test operation example, the composition of components, operation conditions, and the like in the operation are extremely specific values, the present invention is not limited to these values. Further, the composition of the components in the system may be affected by hydrogen, oxygen, or the like to cause a reaction, and the composition may be slightly changed. Accordingly, the numerical values of the examples shown in the table indicate numerical values at a certain time point at the time of implementation.

Comparative example 1

The following experiment was performed in a pilot plant (pilot plant) for acetic acid by methanol process (see fig. 1).

In a reaction vessel (total pressure 2.8MPa (absolute pressure), carbon monoxide partial pressure 1.4MPa (absolute pressure), hydrogen partial pressure 0.04MPa (absolute pressure), reaction temperature 187 deg.C)]The reaction mixture [ composition: methyl iodide (MeI) 7.8%, methyl Acetate (MA) 2.1%, water (H) ₂ O) 2.5%, rhodium complex 910ppm (in terms of Rh), lithium iodide (LiI) 14.1%, acetaldehyde (AD) 250ppm, crotonaldehyde (CR) 1.4ppm, 2-ethylcrotonaldehyde (2 ECR) 1.6ppm, propionic Acid (PA) 250ppm, formic Acid (FA) 40ppm, butyl Acetate (BA) 4.8ppm, residual acetic acid (but containing trace impurities)]400 parts of the mixture was fed to an evaporation tank, and the evaporation tank was heated to evaporate 25% (evaporation rate 25%). The vapor of the evaporation tank (composition: methyl iodide 28.0%, methyl acetate 4.9%, water 1.9%, acetaldehyde 651ppm, crotonaldehyde 1.5ppm, 2-ethylcrotonaldehyde 0.23ppm, propionic acid 73ppm, formic acid 85ppm, butyl acetate 0.7ppm, residual acetic acid (but including trace impurities)]100 parts of the feed was fed to a low-boiling column (actual number of plates: 20, feed position: 2 from below), column top pressure: 250kPa (absolute pressure), column top temperature: 140 ℃ C]Condensing the overhead vapor, and using a decanter [ temperature 40 ℃ C., pressure 0.13MPaG ]]After separation into an aqueous phase and an organic phase, a part (11 parts) of the organic phase was sent to a column for removing acetaldehyde [ 80 stages in the actual plate number, 11 stages from the bottom of the feed position, 280kPa (absolute pressure) at the column top, 52 ℃ at the column top temperature]Acetaldehyde is separated and removed to the outside of the system, and the bottom residue (11 parts substantially equivalent to the feed liquid) after removing acetaldehyde is recycled to the reaction And (4) preparing the system. The remainder of the organic phase (41 parts) was directly recycled to the reaction tank. A part of the aqueous phase was refluxed to the low-boiling-point removal column, and the remainder (1.5 parts) was recycled to the reaction tank as a distillate. The reflux ratio of the aqueous phase was 2 by taking the amount of reflux/distillate as the reflux ratio. 3 parts of the residue was taken as a bottom residue from the low-boiling-point removal column and recycled to the reaction system. 65 parts of the extract was taken as a side Stream (SC) from the middle part (4 layers from the bottom) of the low-boiling column and fed to a dehydration column [ 50 layers in the actual plate number, 34 layers from the bottom in the feed position, 295kPa (absolute pressure) at the column top pressure, 150 ℃ at the column top temperature ]]. A part of the overhead vapor condensate of the dehydration column was refluxed (recycled) to the dehydration column, and the remainder (19 parts) was recycled as a distillate to the reaction system. The reflux ratio (reflux amount/distillate amount) of the dehydration column was set to 0.3. As a result, 46 parts of acetic acid was obtained as a bottom residue from the dehydration column. The crotonaldehyde content in the acetic acid product was 0.99ppm, the 2-ethyl crotonaldehyde content was 0.29ppm, and the butyl acetate content was 0.76ppm. The permanganate time (discoloration time) of the product acetic acid was measured and found to be 50 minutes. The results are shown in Table 3.

Example 1

In comparative example 1 described above, the temperature of the decanter was set to-5.2 ℃, after separation into an aqueous phase and an organic phase, a part (11 parts) of the (organic phase) was sent to an acetaldehyde-removing column [ 80 stages in actual plate number, 11 stages from the bottom of the feed position, 280kPa (absolute pressure) in overhead pressure, 52 ℃ in overhead temperature ] based on the acetaldehyde partition rate, acetaldehyde was separated and removed to the outside of the system, and a bottom liquid (11 parts substantially equivalent to the feed liquid) after removing acetaldehyde was recycled to the reaction system. 20 parts of 41 parts of the organic phase [ composition: alkanes 0.3%, acetaldehyde 1300ppm, methyl acetate 12.5%, water 0.7%, acetic acid 1.9%, crotonaldehyde 1.7ppm, 2-ethyl crotonaldehyde 0.1ppm, butyl acetate 0.3ppm, residual methyl iodide (but containing minor impurities) ] directly recycled to the reaction tank were fed to the decrotonal column [ packed column; the same experiment as in comparative example 1 was carried out except that the number of theoretical plates was 10, the feed position was 5 layers on the basis of theoretical plates from below, the column top pressure was 280kPa (absolute pressure), the column top temperature was 52 ℃ ] (21 parts of the organic phase was directly recycled to the reaction tank), 19.48 parts of [ distillate composition: acetaldehyde 1305ppm, methyl acetate 12.5%, water 0.7%, acetic acid 0.1%, crotonaldehyde 1.5ppm, 2-ethylcrotonaldehyde 0.05ppm, butyl acetate 0.2ppm, residual methyl iodide (but including trace impurities) ] was distilled out at a reflux ratio of 0.01, this was recycled to the decanter, and column bottom raffinate [ column bottom composition: methyl acetate 2.1%, water 1.5%, methyl iodide 5%, crotonaldehyde 6.5ppm, 2-ethylcrotonaldehyde 13.3ppm, butyl acetate 6.9ppm, alkanes 1.2%, residual acetic acid (but including trace impurities) ]0.52 parts was extracted from the column bottom. The composition of each process liquid was changed by this change. As a result, the crotonaldehyde content in the acetic acid product obtained as a column bottom residual liquid from the bottom of the dehydration column was 0.92ppm, the 2-ethylcrotonaldehyde content was 0.28ppm, and the butyl acetate content was 0.71ppm. The permanganate time (discoloration time) of the product acetic acid was measured to be 70 minutes. The results are shown in Table 3.

Example 2

The same experiment as in example 1 was performed except that the temperature of the decanter was changed to 11.0 ℃, the reflux ratio of the low-boiling point removal column was changed to 15, and the reflux ratio of the dehydration column was changed to 10. The composition of each process liquid was changed by this change. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained from the bottom of the dehydration column were 0.23ppm, 0.56ppm, and 2.1ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured and found to be 200 minutes. The results are shown in Table 3.

Example 3

The same experiment as in example 2 was carried out, except that the temperature of the decanter was changed to 22 ℃ and the reflux ratio of the crotonaldehyde column was changed to 0.1. The composition of each process liquid was changed by this change. The distillate amount and the bottoms amount of the crotonaldehyde removal column were changed to 19.52 parts and 0.48 part, respectively. This is a result of the increased separation of water distilled overhead, which is separated from the bottoms. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content and the butyl acetate content of the acetic acid product obtained from the bottom of the dehydration column were 0.21ppm, 0.55ppm and 2ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured to be 220 minutes. The results are shown in Table 3.

Example 4

The same experiment as in example 2 was carried out, except that the temperature of the decanter was set to 29 ℃ and the reflux ratio of the crotonaldehyde column was changed to 1. The composition of each process liquid was changed by this change. The amount of distillate from the crotonaldehyde removing column and the amount of bottoms in the column were changed to 19.56 parts and 0.44 part, respectively. This is a result of further improvement in separation of water distilled at the top of the column and separation from the bottom residue. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained from the bottom of the dehydration column were 0.19ppm, 0.42ppm, and 1.8ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was 280 minutes. The results are shown in Table 3.

Example 5

The same experiment as in example 2 was performed except that the temperature of the decanter was changed to 40.8 ℃ and the reflux ratio of the crotonaldehyde column was changed to 10. The composition of each process liquid was changed by this change. The distillate amount and the bottoms amount of the crotonaldehyde removal column were changed to 19.6 parts and 0.4 part, respectively. This is a result of further improvement in separation of water distilled at the top of the column and separation of the residue from the bottom of the column. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained from the bottom of the dehydration column were 0.15ppm, 0.28ppm, and 1.8ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured and found to be 360 minutes. The results are shown in Table 3.

Example 6

The same experiment as in example 2 was carried out except that the temperature of the decanter was set to 49.9 ℃ and the reflux ratio of the crotonaldehyde removing column was changed to 50. The composition of each process liquid was changed by this change. Further, the amount of distillate in the crotonaldehyde removal column was changed to 19.6 parts and the amount of bottoms in the column was changed to 0.4 part. This is a result of further improvement in separation of water distilled at the top of the column and separation of the residue from the bottom of the column. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content and the butyl acetate content of the acetic acid product obtained from the bottom of the dehydration column were 0.09ppm, 0.14ppm and 1.6ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was 450 minutes. The results are shown in Table 3.

Example 7

The same experiment as in example 5 was carried out except that the temperature of the decanter was set to 40.2 ℃ and the organic phase was selected so that the amount of feed to the acetaldehyde removal column was 11 parts based on the temperature at the time of liquid separation of the condensate at the top of the low-boiling point removal column, and the methyl acetate concentration of the feed to the low-boiling point removal column and the AD distribution ratio were changed. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained as a column bottom residual liquid from the bottom of the dehydration column were 0.3ppm, 0.56ppm, and 2.2ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured to be 150 minutes. The results are shown in Table 4.

Example 8

The same experiment as in example 5 was carried out except that the temperature of the decanter was changed to 40.5 ℃ and the methyl acetate concentration of the feed to the low-boiling point removal column and the methyl acetate concentration of the decanter and the AD distribution ratio were changed. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained as a column bottom residual liquid from the bottom of the dehydration column were 0.22ppm, 0.55ppm, and 2ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured to be 220 minutes. The results are shown in Table 4.

Example 9

An experiment was carried out in the same manner as in example 5 except that the temperature of the decanter was set to 39.9 ℃ and the methyl acetate concentration of the feed to the low-boiling point removal column and the methyl acetate concentration of the decanter and the AD distribution ratio were changed. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content in acetic acid, which was a product obtained as a bottom residue from the bottom of the dehydration column, was 0.18ppm, the 2-ethyl crotonaldehyde content was 0.28ppm, and the butyl acetate content was 1.7ppm. The permanganate time (discoloration time) of the product acetic acid was measured, and found to be 290 minutes. The results are shown in Table 4.

Example 10

The same experiment as in example 5 was carried out except that the temperature of the decanter was changed to 40.7 ℃ and the methyl acetate concentration of the feed to the low-boiling point removal column and the methyl acetate concentration of the decanter and the AD distribution ratio were changed. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content in acetic acid, which was a product obtained as a bottom residue from the bottom of the dehydration column, was 0.12ppm, the 2-ethyl crotonaldehyde content was 0.28ppm, and the butyl acetate content was 1.6ppm. The permanganate time (discoloration time) of the product acetic acid was measured to be 400 minutes. The results are shown in Table 4.

Example 11

The same experiment as in example 5 was carried out except that the temperature of the decanter was changed to 40.1 ℃ and the methyl acetate concentration of the feed to the low-boiling point removal column and the methyl acetate concentration of the decanter and the AD distribution ratio were changed. The composition of each process liquid was changed by this change. As a result, the crotonaldehyde content in acetic acid, which was a product obtained as a bottom residue from the dehydration column, was 0.07ppm, the 2-ethylcrotonaldehyde content was 0.14ppm, and the butyl acetate content was 1.4ppm. The permanganate time (discoloration time) of the product acetic acid was measured to be more than 540 minutes. The results are shown in Table 4.

Example 12

The same experiment as in example 1 was performed except that the temperature of the decanter was changed to 40.1 ℃. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content, the 2-ethyl crotonaldehyde content, and the butyl acetate content in the acetic acid product obtained as a column bottom residual liquid from the bottom of the dehydration column were 0.8ppm, 0.29ppm, and 0.68ppm, respectively. The permanganate time (discoloration time) of the product acetic acid was measured, and found to be 90 minutes. The results are shown in Table 4.

Example 13

An experiment was carried out in the same manner as in example 6 except that the temperature of the decanter was set to 40.3 ℃ and the amount of the aqueous phase to be supplied to the acetaldehyde removal column was selected to 11 parts based on the acetaldehyde partition rate of the condensate at the top of the low-boiling point removal column, and the methyl acetate concentration of the feed to the low-boiling point removal column and the methyl acetate concentration and AD partition rate of the decanter were changed. Due to this change, the composition of each process liquid was changed. As a result, the crotonaldehyde content in acetic acid, which was a product obtained as a bottom residue from the bottom of the dehydration column, was 0.08ppm, the 2-ethyl crotonaldehyde content was 0.12ppm, and the butyl acetate content was 1.3ppm. The permanganate time (discoloration time) of the product acetic acid was 530 minutes. The results are shown in Table 4.

In Table 3, "C _CR "indicates the crotonaldehyde concentration" C _ECR "represents the concentration of 2-ethylcrotonaldehyde" C _BA "indicates the butyl acetate concentration," AD "indicates acetaldehyde," MeI "indicates methyl iodide," MA "indicates methyl acetate," AC "indicates acetic acid. In addition, the numerical values in the columns of the respective components indicate concentrations. Distillate [ CR ]]Feed solution [ CR]The ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate to the crotonaldehyde concentration (mass ppm) in the feed liquid is shown. Tower still residual liquid [ CR]Feed solution [ CR]The ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the bottom liquid to the crotonaldehyde concentration (mass ppm) in the feed liquid is shown.

[ Table 3]

TABLE 3

(remark 1): indicating that no measurement was made.

(remark 2): this step is not performed.

The total value of less than 100[ wt% ] is caused by the influence of rounding up or down the significant figures.

[ Table 4]

TABLE 4

The total value of less than 100[ wt% ] is caused by the influence of rounding-off of the significant figures or less.

[ examination of results ]

Comparative example 1 is an example in which the target (aqueous phase or organic phase) to be subjected to the AD removal treatment is not selected based on the temperature at the time of liquid separation, the AD distribution ratio, and the like, and a part of the organic phase is simply selected and treated, and the crotonaldehyde removal column is not operated.

As is clear from comparison between comparative example 1 and example 1, if the object to be subjected to the AD removing treatment is selected based on the AD distribution ratio in the liquid separation step or the like so that the reflux ratio in the low boiling point removing column becomes a specific value or more and the crotonaldehyde removing column is operated under specific conditions, the crotonaldehyde concentration in the acetic acid product decreases and the discoloration time of the product increases (by 1.4 times). This can be considered as: the object of the AD removal treatment is selected based on the AD distribution ratio in the liquid separation step, acetaldehyde as a by-product is efficiently separated and removed, the reflux ratio in the low-boiling-point removal column is set to a specific value or more, and the crotonaldehyde removal column is operated under specific conditions, and these components exert a synergistic effect.

From the comparison between example 1 and example 2, it is understood that even under the same operating conditions of the crotonaldehyde removing column, the discoloration time of the product is significantly increased (about 3 times) when the reflux ratio of the low boiling point removing column and the dehydrating column is increased. According to examples 2 to 6, even when the reflux ratio of the low-boiling-point removal column was increased and crotonaldehyde was further concentrated at the bottom of the crotonaldehyde removal column under the same conditions as the reflux ratio of the low-boiling-point removal column, the amount of crotonaldehyde discarded increased and the discoloration time of the acetic acid product was increased.

In examples 1 and 2, the reflux ratio of the crotonaldehyde removing column was 0.01, which is extremely low as compared with the usual distillation, but since the distillation almost completely distilled the feed liquid, the distilled amount was almost equal to the feed liquid amount, and was very large, and even a low reflux amount contributed to the separation of crotonaldehyde. In example 6, the methyl iodide concentration in the bottom residue was reduced to the limit, and the discharge of useful methyl iodide was suppressed.

Further, from the results of examples 1 to 6, it is understood that under the condition that the AD concentration and MA concentration in the feed liquid to the low boiling point removal column are constant, the distribution ratio of AD is smaller as the temperature at the time of liquid separation is higher, and even in the system in which methyl acetate is present, the distribution ratio of AD to the organic phase tends to be relatively higher as the temperature is higher. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the temperature at the time of liquid separation is high, the advantage of performing the AD removal treatment on the organic phase is large, and when the temperature at the time of liquid separation is low, the advantage of performing the AD removal treatment on the organic phase is small. On the other hand, even in the system in which methyl acetate is present, when the temperature at the time of liquid separation is low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, and when the temperature at the time of liquid separation is high, the advantage of subjecting the aqueous phase to the AD removal treatment is small. From other points of view, it is preferable to set the liquid temperature at the time of liquid separation in the liquid separation step to a high value when a step or a facility for subjecting the organic phase to the AD removal treatment is used regardless of the presence or absence of methyl acetate. In addition, when a step or an apparatus for subjecting the aqueous phase to the AD removal treatment is used, the liquid temperature at the time of liquid separation in the liquid separation step is preferably set to a low value.

From the results of examples 7 to 11, it is found that the following tendency is present: the distribution ratio of AD to the organic phase is relatively higher as the MA concentration in the feed liquid of the low-boiling removal column is higher and the MA concentrations of the aqueous phase and the organic phase are higher. That is, from the viewpoint of the removal efficiency of acetaldehyde, when the MA concentration of the aqueous phase or the organic phase is high, the advantage of the AD removal treatment of the organic phase is large, and when the MA concentration of the aqueous phase or the organic phase is low, the advantage of the AD removal treatment of the organic phase is small. On the other hand, when the MA concentration of the aqueous phase or the organic phase is low, the advantage of subjecting the aqueous phase to the AD removal treatment is large, whereas when the MA concentration of the aqueous phase or the organic phase is high, the advantage of subjecting the aqueous phase to the AD removal treatment is small. In the case of using a step or an apparatus for subjecting the organic phase to the AD removal treatment, it is preferable to increase the MA concentration of the aqueous phase and/or the organic phase in the liquid separation step by, for example, adjusting the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, or the like. When the step or the facility for subjecting the aqueous phase to the AD removal treatment is used, it is preferable to reduce the MA concentration of the aqueous phase and/or the organic phase in the liquid separation step by adjusting, for example, the reaction conditions in the reaction tank, the evaporation conditions in the evaporator, the distillation conditions of the vapor stream obtained in the evaporation step, and the like. It is understood from the results of examples 7 to 11 that the higher the MA concentration in the feed liquid and the higher the MA concentration in the aqueous phase and the organic phase, the smaller the AD distribution ratio and the MA distribution ratio.

In examples 5 and 7 to 11, the hydrogen partial pressure in the reaction tank, the reflux ratio in the low boiling point removal column, the reflux ratio in the dehydration column, the reflux ratio in the crotonaldehyde removal column, the liquid temperature at the time of liquid separation, and the operating conditions for the AD removal treatment were common, and as a result, the product discoloration time of the 2 nd acetic acid stream was 150 minutes at the minimum (example 7) and 540 minutes at the maximum (example 11), and a difference of about 4 times was exhibited. Here, if attention is paid to the AD distribution ratio, example 7 in which the product discoloration time is the smallest is 4.3, while example 11 in which the product discoloration time is the largest is 1.1. The operating conditions differ in that the MA concentration of the feed to the low-boiling column is varied. This causes a change in MA concentration in the decanter (aqueous phase and organic phase), and a change in AD distribution rate. It is found that, since it is advantageous to subject the aqueous phase to the AD removing treatment when the AD distribution ratio is high to a certain degree (when acetaldehyde is distributed in a large amount in the aqueous phase), and it is advantageous to subject the organic phase to the AD removing treatment when the AD distribution ratio is low to a certain degree (when acetaldehyde is distributed in a large amount in the organic phase), by effectively utilizing the AD distribution ratio in the liquid separation step as an index for performing the AD separation and removal step, it is possible to appropriately select and treat the target to be subjected to the AD removing treatment from at least a part of the aqueous phase and/or the organic phase, and therefore, it is possible to more effectively separate and remove AD as a by-product in an industrially advantageous manner. It is also known that this contributes to an increase in the discoloration time of the product.

Further, from the comparison between example 1 and example 12, it is understood that if the temperature of the decanter is increased and the AD concentration in the organic phase of the decanter is increased, the AD removal amount is increased and the crotonaldehyde production amount is decreased, and therefore the discoloration time of the product is increased to about 1.3 times (90 minutes). This indicates that control of the temperature condition at the time of decanter (liquid temperature at the time of separation) is one of the effective parameters in the AD removal treatment with good efficiency.

As is clear from comparison of examples 11 and 13, when the organic phase and the aqueous phase having substantially the same AD concentration were subjected to AD removal treatment, the crotonaldehyde and 2-ethylcrotonaldehyde concentrations and the discoloration time of the acetic acid product were substantially the same. That is, it is found that, in the case of the AD concentration in the intermediate region, either the aqueous phase or the organic phase may be subjected to the AD removing treatment. In examples 11 and 13, the hydrogen partial pressure in the reaction tank, the reflux ratio in the low-boiling-point removal column, and the reflux ratio in the dehydration column were common, and the reflux ratio was only that in the crotonaldehyde removal column, and example 13 was 5 times as large (reflux ratio 50). However, there was almost no difference in discoloration time of the product, which is considered to be because the purity of the product acetic acid was extremely high, and therefore the effect of the reflux ratio of the crotonaldehyde column hardly exhibited a difference between the case of the reflux ratio of 10 (example 11) and the reflux ratio of 50 (example 13). Further, it was revealed that if the reflux ratio of the crotonaldehyde removing column is increased, the effect is not changed even if the reflux ratio is increased infinitely because the crotonaldehyde separation efficiency is close to a constant value.

In summary, it is found that if crotonaldehyde is converted into 2-ethyl crotonaldehyde and butyl acetate (crotonaldehyde + acetaldehyde → 2-ethyl crotonaldehyde, crotonaldehyde + hydrogen → butanol, butanol + acetic acid → butyl acetate) in the reaction tank by increasing the reflux ratio of the low-boiling-point removal column and/or the dehydration column, concentrating crotonaldehyde to the top of the reaction tank by increasing the reflux ratio of the low-boiling-point removal column, and then distilling the top liquid (for example, organic phase) to effectively remove crotonaldehyde, the improvement in quality beyond the intended level can be achieved by utilizing their synergistic effect.

Further, the reflux ratio of the low-boiling-point removal column is set to a specific value or more, and the crotonaldehyde is effectively removed by the crotonaldehyde removing step; (2) The target of the acetaldehyde removal treatment is determined based on the acetaldehyde distribution ratio in the separation step, and high-quality acetic acid having a good potassium permanganate test value can be produced industrially very efficiently.

In summary, the technical solution and its modifications of the present invention are as follows.

[1] A process for the production of acetic acid comprising:

in the manufacturing method of the present invention, the substrate is,

the separation step comprises a step of separating and obtaining a 1 st acetic acid stream enriched in acetic acid and a top stream enriched in low-boiling components compared with the 1 st acetic acid stream by a 1 st distillation column, wherein the reflux ratio of the 1 st distillation column is set such that the reflux ratio of an aqueous phase is 2 or more when only the aqueous phase is refluxed, the reflux ratio of an organic phase is 1 or more when only the organic phase is refluxed, and the total reflux ratio of the aqueous phase and the organic phase is 1.5 or more when both the aqueous phase and the organic phase are refluxed, and the distillation column is operated so as to satisfy at least 1 of the following conditions (a-i) to (a-iii) in the crotonaldehyde removal step:

(a-i) adjusting the reflux ratio of the distillation column to 0.01 or more;

[2] The process for producing acetic acid according to [1], wherein the separation step satisfies at least one of the following conditions (b-i) to (b-v), and at least a part of the aqueous phase is treated in the acetaldehyde separation and removal step,

(b-ii) the temperature at the time of liquid separation is 70 ℃ or lower;

[3] The process for producing acetic acid according to [1] or [2], wherein in the (b-i), the acetaldehyde concentration in the aqueous phase is 28.1% by mass or less, and the acetaldehyde concentration in the organic phase is 24.8% by mass or less.

[4] The process for producing acetic acid according to any one of the above [1] to [3], wherein in the (b-i), the concentration of acetaldehyde in the aqueous phase is 0.045 to 28.1 mass% (preferably 0.098 to 10 mass%, more preferably 0.098 to 3.0 mass%, particularly preferably 0.098 to 1.0 mass% (e.g., 0.15 to 0.9 mass%).

[5] The process for producing acetic acid according to any one of the above [1] to [4], wherein the acetaldehyde concentration in the organic phase in the (b-i) is 0.013 to 24.8 mass% (preferably 0.030 to 2.0 mass%, more preferably 0.030 to 0.50 mass%, particularly preferably 0.030 to 0.24 mass%).

[6] The process for producing acetic acid according to any one of the above [1] to [5], wherein when the above (b-i) is satisfied, the acetaldehyde concentration in the liquid to be subjected to the liquid separation is 26.0 mass% or less (for example, 0.026 to 26.0 mass%, preferably 0.057 to 10 mass%, more preferably 0.057 to 3.0 mass%, particularly preferably 0.057 to 1.0 mass% (for example, 0.057 to 0.42 mass%).

[7] The process according to any one of the above [1] to [6], wherein the temperature at the time of liquid separation in the step (b-ii) is from-5 ℃ to 70 ℃ (preferably from-5 ℃ to 60 ℃, more preferably from-5 ℃ to 51 ℃ (for example from-5 ℃ to 45 ℃), and still more preferably from-5 ℃ to 41 ℃ (for example from-5 ℃ to 31 ℃).

[8] The process according to any one of [1] to [7], wherein in the (b-iii), the methyl acetate concentration in the aqueous phase is 12.0% by mass or less and the methyl acetate concentration in the organic phase is 47.6% by mass or less.

[9] The process for producing acetic acid according to any one of the above [1] to [8], wherein in the (b-iii), the methyl acetate concentration in the aqueous phase is 1.2 to 12.0 mass% (preferably 2.0 to 12.0 mass%, more preferably 5.0 to 12.0 mass% (e.g., 6.0 to 12.0 mass%)).

[10] The process for producing acetic acid according to any one of the above [1] to [9], wherein the methyl acetate concentration in the organic phase in the (b-iii) is 2.2 to 47.6 mass% (preferably 5.0 to 42 mass%, more preferably 8.0 to 35 mass% (e.g., 10.0 to 30 mass%).

[11] The process for producing acetic acid according to any one of the above [1] to [10], wherein in the above (b-iii), the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase is 4.2 to 59.6 mass% (preferably 6.0 to 54 mass%, more preferably 8.0 to 54 mass%, further preferably 10.0 to 54 mass%, particularly preferably 14.0 to 47 mass% (e.g., 16.0 to 42 mass%).

[12] The process for producing acetic acid according to any one of the above [1] to [11], wherein when the (b-iii) is satisfied, the concentration of methyl acetate in the liquid to be subjected to the liquid separation is 38.2% by mass or less (for example, 2.0 to 38.2% by mass, preferably 5.0 to 31% by mass, more preferably 8.0 to 25% by mass, and still more preferably 10.0 to 25% by mass).

[13] The process for producing acetic acid according to any one of the above [1] to [12], wherein the acetaldehyde partition rate in the (b-iv) is 1.1 to 8.0 (preferably 1.5 to 6.0, more preferably 1.9 to 5.0).

[14] The process for producing acetic acid according to any one of [1] to [13], wherein the methyl acetate distribution ratio in the (b-v) is 0.26 or more (for example, 0.26 to 0.65, preferably 0.28 or more (for example, 0.28 to 0.60)).

[15] The method for producing acetic acid according to any one of [1] to [14], wherein the liquid separation step satisfies at least the (b-i), (b-ii), and (b-iii) at the same time.

[16] The method for producing acetic acid according to any one of [1] to [15], wherein the liquid separation step satisfies at least the (b-i), (b-ii), (b-iii), and (b-iv) at the same time.

[17] The method for producing acetic acid according to any one of [1] to [16], wherein the liquid separation step satisfies all of the conditions of (b-i), (b-ii), (b-iii), (b-iv), and (b-v).

[18] The process for producing acetic acid according to [1], wherein the liquid separation step satisfies at least one of the following conditions (b '-i) to (b' -v), and at least a part of the organic phase is treated in the acetaldehyde separation and removal step,

(b' -ii) the temperature for liquid separation is-5 ℃ or higher;

(b' -v) a methyl acetate partition rate [ { methyl acetate concentration (mass%) of the aqueous phase }/{ methyl acetate concentration (mass%) of the organic phase } ] of 0.8 or less.

[19] The process for producing acetic acid according to item [18], wherein in the (b' -i), the acetaldehyde concentration in the aqueous phase is 0.045% by mass or more and the acetaldehyde concentration in the organic phase is 0.013% by mass or more.

[20] The process for producing acetic acid according to [18] or [19], wherein in the (b' -i), the acetaldehyde concentration in the aqueous phase is 0.045 to 35 mass% (preferably 0.15 to 10 mass%, more preferably 0.2 to 2.0 mass%).

[21] The process for producing acetic acid according to any one of the above [18] to [20], wherein the acetaldehyde concentration in the organic phase in the (b' -i) is 0.013 to 30% by mass (preferably 0.05 to 5.0% by mass, more preferably 0.1 to 1.0% by mass).

[22] The method for producing acetic acid according to any one of [18] to [21], wherein when the (b' -i) is satisfied, the concentration of acetaldehyde in the liquid to be subjected to the liquid separation is 0.026% by mass or more (for example, 0.026 to 32% by mass, preferably 0.10 to 8.0% by mass, and more preferably 0.15 to 1.8% by mass).

[23] The process for producing acetic acid according to any one of [18] to [22], wherein in the (b' -ii), the temperature at the time of liquid separation is 0 ℃ or more (e.g., 0 to 90 ℃, preferably 10 ℃ or more (e.g., 10 to 90 ℃), more preferably 10 ℃ or more (e.g., 10 to 90 ℃), further preferably 20 ℃ or more (e.g., 25 to 90 ℃), a temperature higher than 30 ℃ (e.g., more than 30 ℃ and 90 ℃ or less), a temperature higher than 35 ℃ (e.g., more than 35 ℃ and 90 ℃ or less), a temperature higher than 40 ℃ (e.g., more than 40 ℃ and 90 ℃ or less), and particularly preferably a temperature higher than 70 ℃ (e.g., more than 70 ℃ and 90 ℃ or less).

[24] The process for producing acetic acid according to any one of the above [18] to [23], wherein in the (b' -iii), the methyl acetate concentration in the aqueous phase is 1.2% by mass or more and the methyl acetate concentration in the organic phase is 2.2% by mass or more.

[25] The process for producing acetic acid according to any one of the above [18] to [24], wherein in the (b' -iii), the concentration of methyl acetate in the aqueous phase is 1.2 to 20% by mass, preferably 2.5 to 18% by mass, more preferably 4.0 to 15% by mass, still more preferably 6.0 to 13% by mass, particularly preferably 7.0 to 12% by mass.

[26] The process for producing acetic acid according to any one of [18] to [25], wherein the methyl acetate concentration in the organic phase in the (b' -iii) is 2.2 to 60 mass% (preferably 5.8 to 48 mass%, more preferably 8.0 to 40 mass%, further preferably 10.0 to 30 mass%, particularly preferably 11.0 to 25 mass%).

[27] The process for producing acetic acid according to any one of the above [18] to [26], wherein in the (b' -iii), the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase is 3.4 to 75 mass% (preferably 8.3 to 60 mass% (e.g., 10.0 to 40 mass%), more preferably 15.0 to 50 mass%, and still more preferably 25.0 to 53 mass%).

[28] The process for producing acetic acid according to any one of the above [18] to [27], wherein when the (b' -iii) is satisfied, the concentration of methyl acetate in the liquid to be subjected to the liquid separation is 2.0 to 50% by mass (preferably 5.0 to 38% by mass, more preferably 8.0 to 35% by mass, still more preferably 10.0 to 32% by mass, and particularly preferably 15 to 31% by mass).

[29] The process for producing acetic acid according to any one of [18] to [28], wherein the acetaldehyde partition rate in the (b' -iv) is 3.35 or less ((e.g., 0.6 to 3.35), more preferably 3 or less (0.7 to 3), still more preferably 2.8 or less (e.g., 0.8 to 2.8), particularly preferably 2.5 or less (e.g., 0.8 to 2.5), particularly preferably 2.3 or less (e.g., 0.9 to 2.3), and particularly preferably 2.0 or less (e.g., 1.0 to 2.0)).

[30] The process for producing acetic acid according to any one of [18] to [29], wherein the (b' -v) has a methyl acetate partition rate of 0.7 or less ((e.g., 0.20 to 0.70), more preferably 0.6 or less (e.g., 0.20 to 0.60), still more preferably 0.44 or less (e.g., 0.20 to 0.44), and particularly preferably less than 0.25 (e.g., 0.20 or more and less than 0.25)).

[31] The process for producing acetic acid according to any one of [18] to [30], wherein the liquid separation step satisfies at least the requirements (b ' -i), (b ' -ii), and (b ' -iii) at the same time.

[32] The method for producing acetic acid according to any one of [18] to [31], wherein the liquid separation step satisfies at least the above (b '-i), (b' -ii), (b '-iii), and (b' -iv) at the same time.

[33] The method for producing acetic acid according to any one of [18] to [32], wherein the liquid separation step satisfies all of the conditions (b ' -i), (b ' -ii), (b ' -iii), (b ' -iv), and (b ' -v).

[34] The process for producing acetic acid according to any one of [1] to [33], wherein the separation step comprises:

[35] The process for producing acetic acid according to [34], wherein the separation step further comprises: a dehydration step of separating the 1 st acetic acid stream into a 2 nd overhead stream enriched in water and a 2 nd acetic acid stream enriched in acetic acid compared to the 1 st acetic acid stream using a 2 nd distillation column.

[36] The process for producing acetic acid according to [35], wherein the reflux ratio in the 2 nd distillation column is controlled to 0.1 or more (particularly 0.3 or more, for example, 0.32 or more, preferably 0.35 or more, more preferably 0.4 or more, still more preferably 1 or more, particularly preferably 2 or more).

[37] The process for producing acetic acid according to [36], wherein the upper limit of the reflux ratio in the 2 nd distillation column is 3000 (preferably 1000, more preferably 100, further preferably about 10).

[38] The process for producing acetic acid according to any one of [35] to [37], wherein the concentration of crotonaldehyde in the 2 nd acetic acid stream is 2.0 mass ppm or less (preferably 1.8 mass ppm or less, more preferably 1.5 mass ppm or less, still more preferably 1.2 mass ppm or less, particularly preferably 0.7 mass ppm or less, and particularly preferably 0.5 mass ppm or less).

[39] The process for producing acetic acid according to any one of [35] to [38], wherein the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream is 3.0 mass ppm or less (preferably 1.8 mass ppm or less, more preferably 1.5 mass ppm or less, further preferably 1.2 mass ppm or less, particularly preferably 0.7 mass ppm or less, particularly preferably 0.5 mass ppm or less).

[40]Above-mentioned [35]～[39]The process for producing acetic acid according to any one of the above processes, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream _CR (mass ppm) and 2-Ethylcrotonaldehyde concentration C _ECR (mass ppm) ratio (C) _CR /C _ECR ) Is 50 or less (preferably 35 or less, more preferably 25 or less, further preferably 20 or less, particularly preferably 15 or less).

[41] The process for producing acetic acid according to any one of [35] to [40], wherein the concentration of butyl acetate in the 2 nd acetic acid stream is 15 mass ppm or less (preferably 12 mass ppm or less, more preferably 10 mass ppm or less, further preferably 8 mass ppm or less).

[42]Above [35]]～[41]The process for producing acetic acid according to any one of the above processes, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream _CR (mass ppm) with butyl acetate concentration C _BA (mass ppm) ratio (C) _CR /C _BA ) Is 2.0 or less (preferably 1.5 or less)Lower, more preferably 1.0 or less, and further preferably 0.6 or less).

[43] The process according to any one of [1] to [42], wherein the catalyst system further contains an ionic iodide.

[44] The process for producing acetic acid according to any one of the above [1] to [43], wherein at least a part of a residual liquid after separating acetaldehyde from at least a part of the aqueous phase and/or the organic phase is recycled to the reaction tank.

[45] The process for producing acetic acid according to any one of the above [1] to [44], wherein the hydrogen partial pressure in the reaction vessel is, for example, 0.001MPa (absolute pressure) or more (preferably, 0.005MPa or more, more preferably, 0.01MPa (absolute pressure) or more [ e.g., 0.015MPa or more ], further preferably, 0.02MPa (absolute pressure) or more, particularly preferably, 0.04MPa (absolute pressure) or more, and particularly preferably, 0.06MPa (absolute pressure) or more [ e.g., 0.07MPa (absolute pressure) or more ]).

[46] The process for producing acetic acid according to any one of the above [1] to [45], wherein the acetaldehyde concentration in the reaction mixture in the reaction tank is 500 mass ppm or less (preferably 450 mass ppm or less, more preferably 400 mass ppm or less, further preferably 350 mass ppm or less, particularly preferably 300 mass ppm or less, and particularly preferably 250 mass ppm or less).

[47] The process for producing acetic acid according to any one of [35] to [46], wherein the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.98 mass ppm or less (preferably 0.80 mass ppm or less, more preferably 0.50 mass ppm or less, further preferably 0.30 mass ppm or less) and/or the 2-ethylcrotonaldehyde concentration is 1.0 mass ppm or less (preferably 0.50 mass ppm or less, more preferably 0.30 mass ppm or less, further preferably 0.20 mass ppm or less) and/or the butyl acetate concentration is 15 mass ppm or less (preferably 10 mass ppm or less, more preferably 8 mass ppm or less, further preferably 5 mass ppm or less, particularly preferably 3 mass ppm or less).

[48] The process for producing acetic acid according to any one of [35] to [47], wherein the potassium permanganate test value of the 2 nd acetic acid stream exceeds 50 minutes (preferably 60 minutes or more, more preferably 100 minutes or more, further preferably 120 minutes or more, particularly preferably 180 minutes or more, particularly preferably 240 minutes or more, and further preferably 360 minutes or more).

[49] The process for producing acetic acid according to any one of [34] to [48], wherein the vapor stream supplied to the 1 st distillation column has a crotonaldehyde concentration of 0 to 5.0 mass ppm (preferably 0.01 to 4.0 mass ppm, more preferably 0.1 to 3.0 mass ppm, further preferably 0.2 to 2.0 mass ppm) and/or a 2-ethylcrotonaldehyde concentration of 0 to 3.0 mass ppm (preferably 0.01 to 2.5 mass ppm, more preferably 0.02 to 2.0 mass ppm, further preferably 0.03 to 0.8 mass ppm) and/or a butyl acetate concentration of 0.1 to 13.0 mass ppm (preferably 0.2 to 12.0 mass ppm, more preferably 0.3 to 9.0 mass ppm).

[50] The process for producing acetic acid according to any one of the above [1] to [49], wherein the crotonaldehyde concentration in the 1 st acetic acid stream is 1.3 mass ppm or less (preferably 1.0 mass ppm or less, more preferably 0.85 mass ppm or less, further preferably 0.5 mass ppm or less, and particularly preferably 0.25 mass ppm or less) and/or the 2-ethylcrotonaldehyde concentration is 1.0 mass ppm or less (preferably 0.50 mass ppm or less) and/or the butyl acetate concentration is 15 mass ppm or less (preferably 10 mass ppm or less, more preferably 8 mass ppm or less, further preferably 5 mass ppm or less, and particularly preferably 3 mass ppm or less).

[51] The process for producing acetic acid according to any one of the above [1] to [50], wherein the concentration of crotonaldehyde in the feed liquid to the distillation column in the crotonaldehyde removing step is 0.01 to 50 mass ppm (preferably 0.1 to 50 mass ppm, more preferably 0.3 to 30 mass ppm, further preferably 0.5 to 10 mass ppm, particularly preferably 0.8 to 7.0 mass ppm, particularly preferably 1.0 to 5.0 mass ppm).

[52] The process according to any one of [1] to [51], wherein the crotonaldehyde removal step is performed by operating a distillation column so as to satisfy all of the conditions (a-i) to (a-iii).

[53] The process according to any one of [1] to [52], wherein the crotonaldehyde removing step comprises a batch distillation step.

[54] The process for producing acetic acid according to any one of the above [1] to [53], wherein the amount of the distillation column to be treated in the crotonaldehyde removing step is 0.0001 to 50 parts by mass (preferably 0.001 to 30 parts by mass, more preferably 0.01 to 10 parts by mass, still more preferably 0.1 to 5 parts by mass) per 100 parts by mass of the amount of the vapor stream supplied to the 1 st distillation column.

[55] The process according to any one of [1] to [54], wherein an overhead condensate of the distillation column in the crotonaldehyde removal step is recycled to the aqueous phase and/or the organic phase and/or the reaction tank.

[56] The process for producing acetic acid according to any one of [1] to [55], wherein the reflux ratio of the aqueous phase in the 1 st distillation column is 3 or more (preferably 5 or more, more preferably 8 or more, and further preferably 12 or more) when only the aqueous phase is refluxed in the 1 st distillation column.

[57] The process for producing acetic acid according to any one of [1] to [56], wherein the reflux ratio of the organic phase in the 1 st distillation column is 1.5 or more (preferably 2 or more, more preferably 4 or more, and still more preferably 5 or more) when only the organic phase is refluxed in the 1 st distillation column.

[58] [1] the process for producing acetic acid according to any one of [1] to [57], wherein the reflux ratio in the 1 st distillation column is set such that the total reflux ratio of the aqueous phase and the organic phase is 2.3 or more (preferably 3.5 or more, more preferably 6 or more, and still more preferably 8.5 or more) when both the aqueous phase and the organic phase are refluxed in the 1 st distillation column.

[59] The process for producing acetic acid according to any one of the above [1] to [58], wherein the upper limit of the reflux ratio of the 1 st distillation column is 3000 (preferably 1000, more preferably 100, further preferably 30).

[60] The process for producing acetic acid according to any one of the above [1] to [59], wherein the reflux ratio of the distillation column in the (a-i) is 0.05 or more (preferably 0.5 or more, more preferably 5 or more, further preferably 20 or more, particularly preferably 30 or more).

[61] The process for producing acetic acid according to any one of the above [1] to [60], wherein in the (a-i), the upper limit of the reflux ratio of the distillation column is 1000.

[62] The process for producing acetic acid according to any one of the above [1] to [61], wherein in the (a-ii), the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the distillate of the distillation column to the crotonaldehyde concentration (mass ppm) in the feed liquid is 0.95 or less (preferably 0.80 or less, more preferably 0.70 or less, further preferably 0.60 or less, particularly preferably 0.50 or less, particularly preferably 0.30 or less, and particularly preferably 0.20 or less).

[63] The process for producing acetic acid according to any one of the above [1] to [62], wherein in the (a-iii), the ratio (former/latter) of the crotonaldehyde concentration (mass ppm) in the bottom residue of the distillation column to the crotonaldehyde concentration (mass ppm) in the feed liquid is 1.2 or more (preferably 1.5 or more, more preferably 2.0 or more, further preferably 3.0 or more, particularly preferably 4.0 or more, particularly 5.0 or more, particularly preferably 10 or more, particularly preferably 20 or more).

[64] The process according to any one of [1] to [63], wherein crotonaldehyde is concentrated at a bottom of a distillation column in the crotonaldehyde removal step and discharged out of the system together with acetic acid as a column bottom residue.

[65] The process for producing acetic acid according to any one of the above [1] to [64], wherein the upper limit of the hydrogen partial pressure in the reaction tank is 0.5MPa (absolute pressure) (preferably 0.2MPa (absolute pressure)).

[66] The process for producing acetic acid according to any one of the above [1] to [65], wherein in the crotonaldehyde removing step, the feed position of the feed liquid to the distillation column in the continuous distillation is 20% to 80% (2/10 to 8/10) below the height direction of the distillation column from above.

[67] The process according to any one of [1] to [66], wherein in the crotonaldehyde removal step, at least a part of a condensate of the overhead vapor of the distillation column is refluxed to the distillation column, and at least another part of the condensate is extracted as a distillate and recycled to the aqueous phase and/or the organic phase and/or the reaction tank.

[68] The process according to any one of [1] to [67], wherein a column bottom residual liquid containing crotonaldehyde is extracted from a column bottom of the distillation column in the crotonaldehyde removing step.

[69] The process for producing acetic acid according to any one of [1] to [68], wherein the concentration of crotonaldehyde in the reaction mixture in the reaction tank is 5 mass ppm or less (preferably 3 mass ppm or less, more preferably 2 mass ppm or less).

[70] The process for producing acetic acid according to any one of the above [1] to [69], wherein the concentration of 2-ethylcrotonaldehyde in the reaction mixture in the reaction tank is 5 mass ppm or less (preferably 3 mass ppm or less, more preferably 2 mass ppm or less).

[71] The process for producing acetic acid according to any one of the above [1] to [70], wherein the concentration of butyl acetate in the reaction mixture in the reaction tank is 0.1 to 15 mass ppm (preferably 1 to 12 mass ppm, more preferably 2 to 9 mass ppm).

[72] The process for producing acetic acid according to any one of the above [34] to [71], wherein the concentration of crotonaldehyde in the vapor stream is 0 to 5 mass ppm (preferably 0.1 to 3 mass ppm, more preferably 0.2 to 2 mass ppm).

[73] The process for producing acetic acid according to any one of [34] to [72], wherein the concentration of 2-ethylcrotonaldehyde in the vapor stream is 0 to 3 mass ppm (preferably 0.02 to 2 mass ppm, more preferably 0.03 to 0.8 mass ppm).

[74] The process for producing acetic acid according to any one of the above [34] to [73], wherein the concentration of butyl acetate in the vapor stream is from 0.1 to 13 mass ppm (preferably from 0.2 to 12 mass ppm, more preferably from 0.3 to 9 mass ppm).

[75] The process for producing acetic acid according to any one of [1] to [74], which further comprises: a washing step of separating an off-gas from the process into a stream enriched in carbon monoxide and a stream enriched in acetic acid by subjecting the off-gas to an absorption treatment with an absorption solvent containing at least acetic acid.

Industrial applicability

The method for producing acetic acid of the present invention can be utilized as an industrial production method of acetic acid by a methanol carbonylation process (methanol acetic acid process).

Description of the symbols

1. Reaction tank

2. Evaporation trough

3. 5, 6, 10 distillation column

4. Decanter

7. Ion exchange resin tower

8. Washing system

9. Acetaldehyde separation and removal system

16. Reaction mixture supply line

17. Vapor stream discharge line

18. 19 raffinate recycle line

54. Carbon monoxide-containing gas introduction line

55. 56 potassium hydroxide lead-in circuit

57. Catalyst circulating pump

91. Distillation column (1 st acetaldehyde removing column)

92. Extraction tower

93. Distillation column (No. 2 acetaldehyde removing column)

94. Distillation column (extractive distillation column)

95. Decanter

96. Decanter

97. Distillation column (acetaldehyde removing tower)

98. Distillation column (extractive distillation column)

99. Decanter

200. Chimney type tower tray

300. Liquid-liquid balance measuring device

Claims

1. A process for the production of acetic acid comprising:

a separation step of separating and obtaining a stream containing a metal catalyst, an acetic acid stream enriched in acetic acid, and a stream enriched in low-boiling components compared with the acetic acid stream from the reaction mixture obtained in the carbonylation reaction step, using 1 or more evaporation tanks and/or 1 or more distillation columns;

in the manufacturing method of the present invention, the substrate is,

(a-i) adjusting the reflux ratio of the distillation column to 0.01 or more;

(a-ii) making the ratio of the crotonaldehyde concentration in mass ppm in the distillate to the crotonaldehyde concentration in mass ppm in the feed liquid of the distillation column lower than 1;

(a-iii) making the ratio of the crotonaldehyde concentration in mass ppm in the bottom residue of the distillation column to the crotonaldehyde concentration in mass ppm in the feed liquid greater than 1.

2. The method of producing acetic acid according to claim 1, wherein the liquid separation step satisfies at least one of the following conditions (b-i) to (b-v), and at least a part of the aqueous phase is treated in the acetaldehyde separation and removal step,

(b-ii) the temperature for liquid separation is 70 ℃ or lower;

(b-iv) an acetaldehyde partition rate, that is, { acetaldehyde concentration in mass% in the aqueous phase }/{ acetaldehyde concentration in mass% in the organic phase } of 1.1 or more; and

(b-v) the methyl acetate partition rate, i.e., { methyl acetate concentration in mass% of the aqueous phase }/{ methyl acetate concentration in mass% of the organic phase }, is 0.25 or more.

3. The process for producing acetic acid according to claim 2, wherein in the (b-i), the acetaldehyde concentration in the aqueous phase is 28.1% by mass or less, and the acetaldehyde concentration in the organic phase is 24.8% by mass or less.

4. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the aqueous phase in the (b-i) is 0.045 to 28.1 mass%.

5. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the aqueous phase in the (b-i) is 0.098 to 10% by mass.

6. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the aqueous phase in the (b-i) is from 0.098 to 3.0 mass%.

7. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the organic phase in the (b-i) is 0.013 to 24.8% by mass.

8. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the organic phase in the (b-i) is from 0.030 to 2.0 mass%.

9. The process for producing acetic acid according to claim 2, wherein the concentration of acetaldehyde in the organic phase in the (b-i) is 0.030 to 0.50 mass%.

10. The method for producing acetic acid according to claim 2, wherein when the condition (b-i) is satisfied, the acetaldehyde concentration in the liquid supplied for the liquid separation is 26.0 mass% or less.

11. The process for producing acetic acid according to claim 2, wherein the temperature during liquid separation in (b-ii) is from-5 ℃ to 70 ℃.

12. The process for producing acetic acid according to claim 2, wherein the temperature during liquid separation in (b-ii) is from-5 ℃ to 60 ℃.

13. The process for producing acetic acid according to claim 2, wherein the temperature during liquid separation in (b-ii) is from-5 ℃ to 51 ℃.

14. The process for producing acetic acid according to claim 2, wherein in the step (b-iii), the methyl acetate concentration in the aqueous phase is 12.0% by mass or less, and the methyl acetate concentration in the organic phase is 47.6% by mass or less.

15. The process for producing acetic acid according to claim 2, wherein the methyl acetate concentration in the aqueous phase in the step (b-iii) is 1.2 to 12.0% by mass.

16. The process for producing acetic acid according to claim 2, wherein the concentration of methyl acetate in the organic phase in the (b-iii) is 2.2 to 47.6% by mass.

17. The process for producing acetic acid according to claim 2, wherein the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase in the (b-iii) is 4.2 to 59.6% by mass.

18. The process for producing acetic acid according to claim 2, wherein in the step (b-iii), the concentration of methyl acetate in the aqueous phase is 5.0 to 12.0% by mass, and/or the concentration of methyl acetate in the organic phase is 8.0 to 47.6% by mass, and/or the sum of the concentration of methyl acetate in the aqueous phase and the concentration of methyl acetate in the organic phase is 14.0 to 59.6% by mass.

19. The method for producing acetic acid according to claim 2, wherein when the (b-iii) is satisfied, the concentration of methyl acetate in the liquid supplied to the liquid separation is 38.2 mass% or less.

20. The process for producing acetic acid according to claim 2, wherein the acetaldehyde partition rate in the (b-iv) is 1.1 to 8.0.

21. The process for producing acetic acid according to claim 2, wherein the acetaldehyde partition rate in the (b-iv) is 1.1 to 5.0.

22. The process for producing acetic acid according to claim 2, wherein the partition rate of methyl acetate in the step (b-v) is 0.26 or more.

23. The process for producing acetic acid according to claim 2, wherein the partition rate of methyl acetate in the step (b-v) is 0.25 to 0.65.

24. The method for producing acetic acid according to claim 2, wherein the liquid separation step satisfies at least the conditions (b-i), (b-ii), and (b-iii) at the same time.

25. The method for producing acetic acid according to claim 2, wherein the liquid separation step satisfies at least the (b-i), (b-iii), and (b-iv) at the same time, and/or satisfies at least the (b-i), (b-ii), and (b-v) at the same time.

26. The method for producing acetic acid according to claim 2, wherein the liquid separation step satisfies at least the (b-i), (b-ii), (b-iii), and (b-iv) at the same time, and/or satisfies at least the (b-i), (b-ii), (b-iii), and (b-v) at the same time.

27. The method for producing acetic acid according to claim 2, wherein the liquid separation step satisfies all of the conditions (b-i), (b-ii), (b-iii), (b-iv), and (b-v).

28. The method of producing acetic acid according to claim 1, wherein the liquid separation step satisfies at least one of the following conditions (b '-i) to (b' -v), and at least a part of the organic phase is treated in the acetaldehyde separation and removal step,

(b' -ii) the temperature for liquid separation is-5 ℃ or higher;

(b' -iv) an acetaldehyde partition rate, that is, { acetaldehyde concentration in mass% of the aqueous phase }/{ acetaldehyde concentration in mass% of the organic phase }, of 4.1 or less; and

(b' -v) the methyl acetate partition rate, i.e., { methyl acetate concentration in mass% of the aqueous phase }/{ methyl acetate concentration in mass% of the organic phase }, is 0.8 or less.

29. The process for producing acetic acid according to claim 28, wherein in the (b' -i), the acetaldehyde concentration in the aqueous phase is 0.045% by mass or more, and the acetaldehyde concentration in the organic phase is 0.013% by mass or more.

30. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the aqueous phase in the (b' -i) is from 0.045 to 35% by mass.

31. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the aqueous phase in the (b' -i) is 0.15 to 10% by mass.

32. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the aqueous phase in the (b' -i) is 0.2 to 2.0% by mass.

33. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the organic phase in the (b' -i) is 0.013 to 30% by mass.

34. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the organic phase in the (b' -i) is from 0.05 to 5.0% by mass.

35. The process for producing acetic acid according to claim 28, wherein the concentration of acetaldehyde in the organic phase in the (b' -i) is 0.1 to 1.0 mass%.

36. The method for producing acetic acid according to claim 28, wherein when the (b' -i) is satisfied, the concentration of acetaldehyde in the liquid to be supplied for liquid separation is 0.026% by mass or more.

37. The method for producing acetic acid according to claim 28, wherein when the (b' -i) is satisfied, the concentration of acetaldehyde in the liquid to be subjected to liquid separation is 0.026 to 32% by mass.

38. The method for producing acetic acid according to claim 28, wherein when the condition (b' -i) is satisfied, the concentration of acetaldehyde in the liquid supplied for liquid separation is 0.10 to 8.0% by mass.

39. The method for producing acetic acid according to claim 28, wherein when the condition (b' -i) is satisfied, the concentration of acetaldehyde in the liquid supplied for liquid separation is 0.15 to 1.8% by mass.

40. The process for producing acetic acid according to claim 28, wherein the temperature for liquid separation in the step (b' -ii) is 0 ℃ or higher.

41. The process for producing acetic acid according to claim 28, wherein the methyl acetate concentration in the aqueous phase in (b' -iii) is 1.2% by mass or more, and the methyl acetate concentration in the organic phase is 2.2% by mass or more.

42. The process for producing acetic acid according to claim 28, wherein the methyl acetate concentration in the aqueous phase in the (b' -iii) is from 1.2 to 20% by mass.

43. The process for producing acetic acid according to claim 28, wherein the methyl acetate concentration in the organic phase in the (b' -iii) is 2.2 to 60% by mass.

44. The process for producing acetic acid according to claim 28, wherein the concentration of methyl acetate in the organic phase in the (b' -iii) is 5.8 to 60% by mass.

45. The process for producing acetic acid according to claim 28, wherein the sum of the methyl acetate concentration in the aqueous phase and the methyl acetate concentration in the organic phase in the (b' -iii) is 3.4 to 75% by mass.

46. The method for producing acetic acid according to claim 28, wherein when the condition (b' -iii) is satisfied, the concentration of methyl acetate in the liquid to be supplied for liquid separation is 2.0 to 50% by mass.

47. The process for producing acetic acid according to claim 28, wherein the acetaldehyde partition rate in the (b' -iv) is 3.35 or less.

48. The process for producing acetic acid according to claim 28, wherein the partition rate of methyl acetate in the (b' -v) is 0.7 or less.

49. The process for producing acetic acid according to claim 28, wherein the partition ratio of methyl acetate in the (b' -v) is 0.20 to 0.70.

50. The process for producing acetic acid according to claim 28, wherein the partition rate of methyl acetate in the (b' -v) is 0.6 or less.

51. The process for producing acetic acid according to claim 28, wherein the partition ratio of methyl acetate in the (b' -v) is 0.20 to 0.60.

52. The method for producing acetic acid according to claim 28, wherein the liquid separation step satisfies at least the conditions (b ' -i), (b ' -ii), and (b ' -iii) at the same time.

53. The method for producing acetic acid according to claim 28, wherein the liquid separation step satisfies at least the requirements (b '-i), (b' -iii), and (b '-iv) at the same time, and/or satisfies at least the requirements (b' -i), (b '-ii), and (b' -v) at the same time.

54. The method for producing acetic acid according to claim 28, wherein the liquid separation step satisfies at least the requirements (b '-i), (b' -ii), (b '-iii), and (b' -iv) at the same time, and/or satisfies at least the requirements (b '-i), (b' -ii), (b '-iii), and (b' -v) at the same time.

55. The method for producing acetic acid according to claim 28, wherein the liquid separation step satisfies all of the conditions (b ' -i), (b ' -ii), (b ' -iii), (b ' -iv), and (b ' -v).

56. The method for producing acetic acid according to claim 1 or 2, wherein the separation step comprises:

and a low-boiling removal step of separating the vapor stream in a 1 st distillation column into a 1 st overhead stream enriched in at least one low-boiling component selected from methyl iodide and acetaldehyde and a 1 st acetic acid stream enriched in acetic acid, and condensing and separating the 1 st overhead stream to obtain an aqueous phase and an organic phase.

57. The method of producing acetic acid according to claim 56, further comprising: a dehydration step of separating the 1 st acetic acid stream into a 2 nd overhead stream enriched in water and a 2 nd acetic acid stream enriched in acetic acid compared to the 1 st acetic acid stream using a 2 nd distillation column.

58. The process for producing acetic acid according to claim 57, wherein the reflux ratio of the 2 nd distillation column is 0.1 or more.

59. The process for producing acetic acid according to claim 57, wherein the reflux ratio of the 2 nd distillation column is 0.3 or more.

60. The process for producing acetic acid according to claim 57, wherein the reflux ratio of the 2 nd distillation column is 1.0 or more.

61. The process for producing acetic acid according to claim 57, wherein the upper limit of the reflux ratio of the 2 nd distillation column is 3000.

62. The process for producing acetic acid according to claim 57, wherein the upper limit of the reflux ratio of the 2 nd distillation column is 100.

63. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.98 mass ppm or less, and/or the 2-ethyl crotonaldehyde concentration is 1.0 mass ppm or less, and/or the butyl acetate concentration is 15 mass ppm or less.

64. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.80 mass ppm or less and/or the 2-ethyl crotonaldehyde concentration is 0.50 mass ppm or less and/or the butyl acetate concentration is 10 mass ppm or less.

65. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.7 ppm by mass or less.

66. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration in the 2 nd acetic acid stream is 0.5 ppm by mass or less.

67. The process for producing acetic acid according to claim 57, wherein the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream is 0.7 ppm by mass or less.

68. The process for producing acetic acid according to claim 57, wherein the concentration of 2-ethylcrotonaldehyde in the 2 nd acetic acid stream is 0.5 ppm by mass or less.

69. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream is in mass ppm _CR With a 2-ethylcrotonaldehyde concentration C in mass ppm _ECR Ratio of C _CR /C _ECR Is 50 or less.

70. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream is mass ppm _CR With a 2-ethylcrotonaldehyde concentration C in ppm by mass _ECR Ratio of C _CR /C _ECR Is 25 or less.

71. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream is mass ppm _CR With a 2-ethylcrotonaldehyde concentration C in mass ppm _ECR Ratio of C _CR /C _ECR Is 15 or less.

72. The process for producing acetic acid according to claim 57, wherein the concentration of butyl acetate in the 2 nd acetic acid stream is 12 ppm by mass or less.

73. The process for producing acetic acid according to claim 57, wherein the concentration of butyl acetate in the 2 nd acetic acid stream is 10 ppm by mass or less.

74. The process for producing acetic acid according to claim 57, wherein the concentration of butyl acetate in the 2 nd acetic acid stream is 8 ppm by mass or less.

75. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream is in mass ppm _CR With a butyl acetate concentration C in mass ppm _BA Ratio of C _CR /C _BA Is 2.0 or less.

76. The process for producing acetic acid according to claim 57, wherein the crotonaldehyde concentration C in the 2 nd acetic acid stream is mass ppm _CR With a butyl acetate concentration C in mass ppm _BA Ratio of C _CR /C _BA Is 1.5 or less.

77. The process for producing acetic acid according to claim 57, wherein the 2 nd acetic acid stream has a potassium permanganate test value of more than 50 minutes.

78. The process for producing acetic acid according to claim 57, wherein the 2 nd acetic acid stream has a potassium permanganate test value of 60 minutes or more.

79. The process for producing acetic acid according to claim 56, wherein the concentration of crotonaldehyde in the vapor stream supplied to the 1 st distillation column is 0 to 5.0 mass ppm, and/or the concentration of 2-ethylcrotonaldehyde is 0 to 3.0 mass ppm, and/or the concentration of butyl acetate is 0.1 to 13.0 mass ppm.

80. The process for producing acetic acid according to claim 56, wherein the vapor stream supplied to the 1 st distillation column has a crotonaldehyde concentration of 0.01 to 4.0 mass ppm, and/or a 2-ethylcrotonaldehyde concentration of 0.01 to 2.5 mass ppm, and/or a butyl acetate concentration of 0.2 to 12.0 mass ppm.

81. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the 1 st acetic acid stream is 1.3 mass ppm or less, and/or the 2-ethyl crotonaldehyde concentration is 1.0 mass ppm or less, and/or the butyl acetate concentration is 15 mass ppm or less.

82. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the 1 st acetic acid stream is 1.0 mass ppm or less and/or the 2-ethyl crotonaldehyde concentration is 0.50 mass ppm or less and/or the butyl acetate concentration is 10 mass ppm or less.

83. The process for the production of acetic acid according to claim 1 or 2, wherein the catalytic system further comprises an ionic iodide.

84. The process for producing acetic acid according to claim 1 or 2, wherein at least a part of a residual liquid after separating acetaldehyde from at least a part of the aqueous phase and/or the organic phase is recycled to the reaction tank.

85. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is 0.01 to 50 ppm by mass.

86. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is 0.1 to 50 ppm by mass.

87. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is 0.3 to 30 ppm by mass.

88. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed solution to the distillation column in the crotonaldehyde removal step is 0.5 to 10 ppm by mass.

89. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed solution to the distillation column in the crotonaldehyde removing step is 0.8 to 7.0 ppm by mass.

90. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the feed liquid to the distillation column in the crotonaldehyde removing step is 1.0 to 5.0 ppm by mass.

91. The process for producing acetic acid according to claim 1 or 2, wherein in the crotonaldehyde removing step, the distillation column is operated so as to satisfy all of the conditions (a-i) to (a-iii) above.

92. The process for producing acetic acid according to claim 1 or 2, wherein in the crotonaldehyde removing step, distillation is performed in a batch process.

93. The process for producing acetic acid according to claim 56, wherein the amount of the distillation column to be treated in the crotonaldehyde removing step is 0.0001 to 50 parts by mass relative to 100 parts by mass of the amount of the vapor stream supplied to the 1 st distillation column.

94. The process for producing acetic acid according to claim 56, wherein the amount of the distillation column to be treated in the crotonaldehyde removing step is 0.001 to 30 parts by mass relative to 100 parts by mass of the amount of the vapor stream supplied to the 1 st distillation column.

95. The method for producing acetic acid according to claim 1 or 2, further comprising: a washing step of separating an off-gas from the process into a stream enriched in carbon monoxide and a stream enriched in acetic acid by subjecting the off-gas to an absorption treatment with an absorption solvent containing at least acetic acid.

96. The process for producing acetic acid according to claim 1 or 2, wherein the hydrogen partial pressure in the reaction vessel is 0.001MPa or more in absolute terms.

97. The process for producing acetic acid according to claim 1 or 2, wherein the acetaldehyde concentration in the reaction mixture in the reaction tank is 500 ppm by mass or less.

98. The process for producing acetic acid according to claim 1 or 2, wherein an overhead condensate of the distillation column in the crotonaldehyde removing step is recycled to the aqueous phase and/or the organic phase and/or the reaction tank.

99. The process for producing acetic acid according to claim 1 or 2, wherein the reflux ratio of the aqueous phase to the 1 st distillation column is 3 or more when only the aqueous phase is refluxed to the 1 st distillation column.

100. The process for producing acetic acid according to claim 1 or 2, wherein the reflux ratio of the organic phase in the 1 st distillation column is 1.5 or more when only the organic phase is refluxed to the 1 st distillation column.

101. The process for producing acetic acid according to claim 1 or 2, wherein the reflux ratio of the 1 st distillation column is 2.3 or more in total of the aqueous phase and the organic phase when both the aqueous phase and the organic phase in the 1 st distillation column are refluxed to the 1 st distillation column.

102. The process for producing acetic acid according to claim 1 or 2, wherein the upper limit of the reflux ratio of the 1 st distillation column is 3000.

103. The process for producing acetic acid according to claim 1 or 2, wherein the reflux ratio of the distillation column in the step (a-i) is 0.05 or more.

104. The process for producing acetic acid according to claim 1 or 2, wherein in the (a-i), the upper limit of the reflux ratio of the distillation column is 1000.

105. The process for producing acetic acid according to claim 1 or 2, wherein in the step (a-ii), the ratio of the crotonaldehyde concentration in mass ppm in the distillate of the distillation column to the crotonaldehyde concentration in mass ppm in the feed liquid is set to 0.95 or less.

106. The process for producing acetic acid according to claim 1 or 2, wherein in the (a-iii), the ratio of the crotonaldehyde concentration in mass ppm in the bottom residue of the distillation column to the crotonaldehyde concentration in mass ppm in the feed liquid is set to 1.2 or more.

107. The process for producing acetic acid according to claim 1 or 2, wherein crotonaldehyde is concentrated at the bottom of the distillation column in the crotonaldehyde removing step, and discharged as a bottom residue together with acetic acid to the outside of the system.

108. The process for producing acetic acid according to claim 1 or 2, wherein the upper limit of the hydrogen partial pressure in the reaction tank is 0.5MPa abs.

109. The process for producing acetic acid according to claim 1 or 2, wherein in the crotonaldehyde removal step, the feed position of the feed liquid to the distillation column in the continuous distillation is from 2/10 to 8/10 of the height direction of the distillation column from the top.

110. The process for producing acetic acid according to claim 1 or 2, wherein in the crotonaldehyde removal step, at least a part of a condensate of the overhead vapor of the distillation column is refluxed to the distillation column, and at least another part of the condensate is extracted as a distillate and recycled to the aqueous phase and/or organic phase and/or the reaction tank.

111. The process for producing acetic acid according to claim 1 or 2, wherein in the crotonaldehyde removal step, a column bottom residual liquid containing crotonaldehyde is extracted from a column bottom of the distillation column.

112. The process for producing acetic acid according to claim 1 or 2, wherein the crotonaldehyde concentration in the reaction mixture in the reaction tank is 5 ppm by mass or less.

113. The process for producing acetic acid according to claim 1 or 2, wherein the concentration of 2-ethylcrotonaldehyde in the reaction mixture liquid in the reaction tank is 5 ppm by mass or less.

114. The process for producing acetic acid according to claim 1 or 2, wherein the concentration of butyl acetate in the reaction mixture in the reaction tank is 0.1 to 15 ppm by mass.

115. The process for producing acetic acid according to claim 56, wherein the crotonaldehyde concentration of the vapor stream is 0 to 5 ppm by mass.

116. The process for producing acetic acid according to claim 56, wherein the vapor stream has a 2-ethylcrotonaldehyde concentration of 0 to 3 ppm by mass.

117. The process for producing acetic acid according to claim 56, wherein the concentration of butyl acetate in the vapor stream is from 0.1 to 13 ppm by mass.